Catalyst method and apparatus for an on-stream particle replacement system for countercurrent contact of a gas and liquid feed stream with a packed bed

ABSTRACT

This invention makes possible substantially continuous flow of uniformly distributed hydrogen and hydrocarbon liquid across a densely packed catalyst bed to fill substantially the entire volume of a reactor vessel by introducing the fluids as alternate annular rings of gas and liquid (i.e. a mixture of liquid hydrocarbon and a hydrogen-containing gas) at a rate insufficient to levitate or ebullate the catalyst bed. Catalyst are selected by density, shape and size at a design feed rate of liquids and gas to prevent ebullation of the packed bed at the design feed rates. Catalysts are selected by measuring bed expansion, such as in a large pilot plant run, with hydrocarbon, hydrogen, and catalyst at the design pressures and flow velocities. The liquid and gas components of the feed flow into the bed in alternate annular rings across the full area of the bed. At the desired flow rate, such catalyst continually flows in a plug-like manner downwardly through the reactor vessel by introducing fresh catalyst at the top of the catalyst bed by laminarly flowing such catalyst in a liquid stream on a periodic or semicontinuous basis. Catalyst is similarly removed by laminarly flowing catalyst particles in a liquid stream out of the bottom of the catalyst bed. Intake for such flow is out of direct contact with the stream of gas flowing through the bed and the flow path is substantially constant in cross-sectional area and greater in diameter by several times than the diameter of the catalyst particles. The catalyst of this invention produces a plug-flowing substantially packed bed of hydroprocessing catalyst which occupies at least about 75% by volume of the reactor volume.

This application is a continuation of U.S. application Ser. No.08/235,043 filed Apr. 29, 1994, abandoned, which is acontinuation-in-part of U.S. application Ser. No. 08/215,254, filed Mar.21, 1994, now U.S. Pat. No. 5,409,598, which is a continuation of U.S.application Ser. No. 08/014,847, filed Feb. 8, 1993, now U.S. Pat. No.5,302,357, which is a continuation of U.S. application Ser. No.07/727,656, filed Jul. 9, 1991, now abandoned, which is a divisional ofU.S. application Ser. No. 07/381,948, filed Jul. 19, 1989, now U.S. Pat.No. 5,076,908. Benefit of the earliest filing date is claimed,especially with respect to all common subject matter.

BACKGROUND OF THE INVENTION

1. Field of the Invention

The present invention relates to a catalyst and to an on-stream catalystreplacement during hydroprocessing of a hydrocarbon feed stream.

More particularly, it relates to a catalyst, a method of, and apparatusfor, economically utilizing space within a hydroprocessing vessel over awide range of processing rates without substantial fluidization orebullation of a packed bed of catalyst during high counterflow rates ofthe hydrocarbon feed and a hydrogen containing gas through the packedbed, while maintaining continuous or intermittent replacement ofcatalyst for plug-like flow of the bed through the vessel. Such plugflow with high processing rates is obtained by selecting the size, shapeand density of the catalyst particles to prevent ebullation and bedexpansion at the design flow rate so as to maximize the amount ofcatalyst in the vessel during normal operation and during catalysttransfer. Catalysts are selected by measuring bed expansion, such as ina large pilot plant run, with hydrocarbon, hydrogen and catalyst at thedesign pressures and flow velocities within the available reactionvolume of the vessel. Catalyst is removed from the bed by laminar flowof the catalyst particles in a liquid slurry system in which the liquidflow line is uniform in diameter, and substantially larger than thecatalyst particles, throughout the flow path between the reactor vesseland a pressurizable vessel including passageways through the flowcontrol valves. The catalyst of the present invention may also beadvantageously practiced in hydrocarbon reactor systems that utilize an"expanded catalyst bed" such as the ebullated beds with catalyst inrandom motion as described in U.S. Pat. No. 4,571,326 and U.S. Pat. No.4,744,887.

2. Description of the Prior Art

Hydroprocessing or hydrotreatment to remove undesirable components fromhydrocarbon feed streams is a well known method of catalyticallytreating such heavy hydrocarbons to increase their commercial value."Heavy" hydrocarbon liquid streams, and particularly reduced crude oils,petroleum residua, tar sand bitumen, shale oil or liquified coal orreclaimed oil, generally contain product contaminants, such as sulfur,and/or nitrogen, metals and organo-metallic compounds which tend todeactivate catalyst particles during contact by the feed stream andhydrogen under hydroprocessing conditions. Such hydroprocessingconditions are normally in the range of 212 degree(s) F. to 1200degree(s) F. (100 degree(s) to 650 degree(s) C.) at pressures of from 20to 300 atmospheres. Generally such hydroprocessing is in the presence ofcatalyst containing group VI or VIII metals such as platinum,molybdenum, tungsten, nickel, cobalt, etc., in combination with variousother metallic element particles of alumina, silica, magnesia and soforth having a high surface to volume ratio. More specifically, catalystutilized for hydrodemetallation, hydrodesulfurization,hydrodenitrification, hydrocracking etc., of heavy oils and the like aregenerally made up of a carrier or base material; such as alumina,silica, silica-alumina, or possibly, crystalline aluminosilicate, withone more promoter(s) or catalytically active metal(s) (or compound(s))plus trace materials. Typical catalytically active metals utilized arecobalt, molybdenum, nickel and tungsten; however, other metals orcompounds could be selected dependent on the application.

Because these reactions must be carried out by contact of ahydrogen-containing gas with the hydrocarbon feed stream at elevatedtemperatures and pressures, the major costs of such processing areessentially investment in vessels and associated furnaces, heatexchangers, pumps, piping and valves capable of such service and thereplacement cost of catalyst contaminated in such service. Commercialhydroprocessing of relatively low cost feed stocks such as reduced crudeoils containing pollutant compounds, requires a flow rate on the orderof a few thousand up to one hundred thousand barrels per day, withconcurrent flow of hydrogen at up to 10,000 standard cubic feet perbarrel of the liquid feed. Vessels capable of containing such a reactionprocess are accordingly cost-intensive both due to the need to containand withstand corrosion and metal embrittlement by the hydrogen andsulfur compounds, while carrying out the desired reactions, such asdemetallation, denitrification, desulfurization, and cracking atelevated pressure and temperatures. For example, because of metallurgyand safety requirements, such vessels may cost on the order of $700.00per cubic foot of catalyst capacity. Thus a vessel capable of handling25,000 barrels per day of a hydrocarbon feed stream may run on the orderof $4,000,000 to $5,000,000. Pumps, piping and valves for handling fluidstreams containing hydrogen at such pressures and temperatures are alsocostly, because at such pressures seals must remain hydrogen imperviousover extended service periods of many months.

Further, hydroprocessing catalyst for such a reactor, which typicallycontains metals such as titanium, cobalt, nickel, tungsten, molybdenum,etc., may involve a catalyst inventory of 500,000 pounds or more at acost of $2 to $4/lb. Accordingly, for economic feasibility in commercialoperations, the process must handle high flow rates and the vesselshould be filled with as much catalyst inventory as possible to maximizecatalyst activity and run length. Additionally, the down-time forreplacement or renewal of catalyst must be as short as possible.Further, the economics of the process will generally depend upon theversatility of the system to handle feed streams of varying amounts ofcontaminants such as sulfur, nitrogen, metals and/or organic-metalliccompounds, such as those found in a wide variety of the more plentiful(and hence cheaper) reduced crude oils, residua, or liquified coal, tarsand bitumen or shale oils, as well as used oils, and the like. Thefollowing three acceptable reactor technologies are currently availableto the industry for hydrogen upgrading of "heavy" hydrocarbon liquidstreams: (i) fixed bed reactor systems; (ii) ebullated or expanded typereactor systems which are capable of onstream catalyst replacement andare presently known to industry under the trademarks H-Oil® andLC-Fining®; and (iii) the substantially packed-bed type reactor systemhaving an onstream catalyst replacement system, as more particularlydescribed in U.S. Pat. No. 5,076,908 to Stangeland et al, having acommon assignee with the current inventions and discoveries.

A fixed bed reactor system may be defined as a reactor system having oneor more reaction zone(s) of stationary catalyst, through which feedstreams of liquid hydrocarbon and hydrogen flow downwardly andconcurrently with respect to each other.

An ebullated or expanded bed reactor system may be defined as a reactorsystem having an upflow type single reaction zone reactor containingcatalyst in random motion in an expanded catalytic bed state, typicallyexpanded from 10% by volume to about 35% or more by volume above a"slumped" catalyst bed condition (e.g. a non-expanded or non-ebullatedstate).

As particularly described in U.S. Pat. No. 5,076,908 to Stangeland etal, the substantially packed-bed type reactor system is an upflow typereactor system including multiple reaction zones of packed catalystparticles having little or no movement during normal operatingconditions of no catalyst addition or withdrawal. In the substantiallypacked-bed type reactor system of Stangeland et al., when catalyst iswithdrawn from the reactor during normal catalyst replacement, thecatalyst flows in a downwardly direction under essentially plug flow orin an essentially plug flow fashion, with a minimum of mixing withcatalyst in layers which are adjacent either above or below the catalystlayer under observation.

Of the three acceptable reactor systems, most hydroconversion reactorsystems presently in operation on a worldwide basis are fixed bedreactors wherein a liquid hydrocarbon feed and a hydrogen stream flowconcurrently through the catalyst beds in a downward flow path. Whilethese fixed bed downflow type processes assure maximum density or volumeof catalyst within a reaction zone without expansion of the bed, theyare limited by the tendency of the catalyst to form local deposits offeed metals and other contaminates, particularly in the top catalyst bed(i.e. first reaction zone), affecting distribution and reaction rates.As reactor average temperatures are progressively increased to maintainprocessing objectives under conditions of increasing local metaldeposits, catalyst deactivation due to carbon deposition accelerates.When processing objectives can no longer be maintained due to catalystdeactivation (i.e. normally recognized as "End of Run" conditions), thereactor system must be taken offstream for catalyst regeneration orreplacement. Accordingly, in general, it is preferred to counterflow thecatalyst and process fluid streams relative to each other. However, asnoted above, when the process feed rates are high, the volume ofcatalyst that can be contained by the vessel may be as little as 10% ofthe original settled volume. At lower fluid velocities, catalyst volumemay be up to about 80% to 90%, but useful reaction space for the processis still wasted and turbulence causes axial mixing of the catalyst.Therefore, one particular object of the present invention is to run acounterflow processing system where the catalyst and fluid velocitycombinations limit bed expansion to less than 10% by length (morepreferably less than about 5% by length, most preferably less than 2% oreven less than 1% by length) beyond a substantially full axial length ofthe bed in a packed bed state.

It is also known to use a series of individual vessels stacked one abovethe other, with fluid flow either co-current or counterflow to catalyst.In such a process, catalyst moves by gravity from the upper vessel to alower vessel by periodically shutting off, or closing, valves betweenthe individual vessels. In a counterflow system, this permits removal ofcatalyst from the lowermost or first stage vessel, where the mostcontaminated, or raw, feed stock, originally contacts the catalyst. Inthis way, most of the major contaminating components in the hydrocarbonstream are removed before the hydrocarbon material reaches majorconversion steps of the process performed in higher vessels of thestacked series. Thus, most of the deactivating components of the feedstream are removed before it reaches the least contaminated catalystadded to the topmost vessel. However, such systems require valvessuitable for closing off catalyst flow against catalyst trapped in theline. Hence, valve life is relatively short and down-time forreplacement or repair of the valves is relatively costly.

Since the late 1960's, there have been several heavy oil hydroprocessingunits built and brought on stream that utilize the ebullated or expandedcatalyst bed reactor technology where a hydrocarbon feed stream andhydrogen gas flow upwardly through a dilute phase reaction zone ofcatalyst in random motion. Stated alternatively, continuous operation ofan ebullated or expanded bed hydroprocessing system include the upwardflow of a hydrocarbon feed stream and hydrogen gas through a singlecatalyst containing vessel or a series of catalyst containing vessels.Reactor liquid is recirculated internally at rates sufficient to expandor ebullate the catalyst to produce a dilute phase reaction zone ofcatalyst in random or ebullating motion. Catalyst is replaced bycontinuous or periodic, onstream removal of catalyst from the vesselfollowed by addition. As noted above, such ebullation tends to increasethe fluid volume in the vessel relative to catalyst volume necessary tohydroprocess the feed stream and hydrogen with the catalyst, withadequate contact time to react the fluids. Further, such ebullated bedstend to result in separation or segregation of "fines" from the larger(and heavier) particles as they pass downwardly through the upflowstreams. As frequently happens, and especially where the catalyst islocally agitated, as by eddy currents, the particles tend to abrade bysuch higher flow rates of the feed streams through the ebullating bed.Depending on the size of the fines, they either travel upward where theycontaminate the product or they tend to accumulate in the reactorbecause they cannot work their way down to the bottom of the bed. Suchcounterflow systems have also been used because of the relative ease ofwithdrawing limited amounts of the ebullated catalyst in a portion ofthe reacting hydrocarbon and hydrogen fluids, particularly where suchturbulent flow of the catalyst is needed to assist gravity drainagethrough a funnel-shaped opening into a central pipe at the bottom of avessel.

While it has been proposed heretofore to use plug-flow or packed-bedflow of catalyst to reduce such agitation and thus assure uniformdisbursement of hydrogen throughout the liquid volume flowing upwardlythrough the catalyst bed, in general such flow has been controlled bylimiting the maximum flow rate that can be tolerated without ebullatingor levitating the bed more than about 10%. Further in prior systemswhere expansion of the bed is limited, hydrogen flow rates are madesufficiently high at the bottom of the bed to assure relative turbulenceof the catalyst at the withdrawal point in the vessel. While this doesassure such turbulence, it also wastes space, damages the catalyst andpermits direct entrainment of hydrogen with catalyst entering thewithdrawal tube. Such turbulent flow of catalyst is apparently necessaryto assist gravity removal of catalyst from the vessel.

The basic process design of the ebullated bed reactors with appropriatemechanical features overcome some of the limitations of the conventionalfixed bed reactor. The ebullated or expanded catalyst bed reactorschemes provide ability to replace catalyst on stream and operate with avery "flat" reaction zone temperature profile instead of the steeperpyramiding profile of conventional fixed bed reactors. The nature of theprocess, with a broad spectrum of catalyst size, shape, particledensity, and activity level in random motion in a "dilute phase reactionzone," creates near isothermal temperature conditions, with only a fewdegrees temperature rise from the bottom to the top of the reactionzone. Quench fluids are not normally required to limit reaction ratesexcept in cases when series reactors are applied. In other words, thereactor internal recycle oil flow, used to expand (or ebullate) thecatalyst bed and maintain distribution (typically 10 to 1 ratio of freshoil feed) serves also as "internal quench" to control reaction rate andpeak operating temperatures. Because the highest temperaturesexperienced in the reactors are only a few degrees above the averagetemperature required to maintain processing objectives and not thehigher "end of run" peak temperatures experienced in fixed bed reactorsystems, the accelerated fouling rate of the catalyst by carbondeposition experienced in conventional fixed bed reactor systems at "endof run" conditions is minimized; however, the normal carbon depositionrate is much greater than that of the fixed bed reactor due to overalloperating conditions.

Unfortunately, implementing the ebullated bed technology results ininefficient use of reactor volume and less than optimum usage ofhydroconversion catalyst. Catalyst replacement rates for ebullated bedreactors are based on maintaining "catalyst equilibrium conditions"necessary to maintain processing objectives. The backmixing nature ofebullated catalyst beds, combined with the characteristics of thetypical extrudate catalyst particulate used (i.e. a full range of sizesand shapes), promote isothermal temperature conditions but createselectivity difficulties in regard to the withdrawal of expendedcatalyst. Fresh or partially expended catalyst commingle with expendedcatalyst withdrawn from the bottom of the catalyst bed requirescomplicated procedures and equipment to recover, and are usuallydiscarded with minimum recovery value. In other words, use of varioussize and shape catalyst in ebullated bed type reactors leads to somewhatinefficient use of catalyst value.

The additional reactor volume required for the ebullated bed process isto accommodate the expansion of the catalyst load by 25-35% of itsoriginal slumped (or "packed bed") volume or height, by controlling thevelocity of an internal liquid recycle stream. Space required within anebullated bed reactor for the disengagement of solids and catalyst bedlevel controls, and the space required to satisfy suction conditions forthe reactor internal recycle pump, combined with the space the pumpsuction line occupies, consumes a substantial amount of space availablewithin the ebullated bed reactor. Additional disadvantages of theebullated bed technology are the added cost, maintenance, and thereliability of a single supply source for the reactor recycle pump whichis required to expand the catalyst bed. In order to compare efficientuse of reactor volume purchased, the following examples are offered.

If ebullating bed reactor technology is implemented and 13-foot diameterreactors are selected, the tangent line to tangent line dimensionsrequired for the 13-foot diameter reactors would be approximately 60feet in order to load approximately 5,000 cubic feet (or about 175,000lb) of typical hydroprocessing catalyst. Thus, the 5,000 cubic feet ofcatalyst occupies about 63% by volume of the approximately 7,900 cubicfeet of reactor volume available between the bottom and top tangent lineof the reactor. In the case of fixed bed reactors, in order to load5,000 cubic feet of typical hydroprocessing catalyst in 13-foot diameterreactors would require tangent line to tangent line dimensions for the13-foot diameter reactors of about 43 feet; however, the operating runlength for the fixed bed reactors would be short as catalyst could notbe replaced on stream. Should the 60-foot tangent line to tangent linedimensions required for the ebullated bed reactors be maintained for afixed bed reactor, an additional catalyst volume of approximately 2000cubic feet could be loaded.

In order to load 5,000 cubic feet of typical hydroprocessing catalyst ina 13-foot diameter bed reactor with the broad features and descriptionsas disclosed in U.S. Pat. No. 5,076,908 to Stangeland et al, wouldrequire tangent line to tangent line dimensions of approximately 41feet. There would be a reduction of reactor empty weight of between 100to 200 tons, depending on the design pressure specification. Should the60-foot tangent line to tangent line dimensions for a Stangeland et alreactor be maintained, an additional catalyst volume of approximately2500 cubic feet could be loaded.

As particularly distinguished from prior known methods of on-streamcatalyst replacement in hydroprocessing, the method and apparatus inU.S. Pat. No. 5,076,908 to Stangeland el al more specifically provides asystem wherein plug flow of the catalyst bed is maintained over a widerange of counterflow rates of a hydrocarbon feed stream and hydrogen gasthroughout the volume of the substantially packed catalyst bed. Suchpacked bed flow maintains substantially maximum volume and density ofcatalyst within a given vessel's design volume by controlling the size,shape and density of the catalyst so that the bed is not substantiallyexpanded at the design rate of fluid flow therethrough. The proper size,shape and density are determined by applying coefficients gained duringextensive studying of bed expansion in a large pilot plant runs withhydrocarbon, hydrogen and catalyst at the design pressures and flowvelocities as particularly described below.

To further control such packed bed flow, the bed level of catalystwithin the vessel is continuously measured, as by gamma ray absorption,to assure that little ebullation of the bed is occurring. Such controlis further promoted by evenly distributing both the hydrogen and liquidfeed throughout the length of the bed by concentrically distributingboth the hydrogen gas component and the hydrocarbon fluid feed componentin alternate, concentric annular paths across the full horizontalcross-sectional area of the vessel as they both enter the catalyst bed.Additionally, and as desirable, hydrogen is evenly redistributed and ifneeded, augmented, through a quench system at one or more intermediatelevels along the length of the catalyst bed. Equalizing hydrogen andliquid feed across the full horizontal area along the length of thepacked particle bed prevents local turbulence and undesirable verticalsegregation of lighter particles from heavier particles flowing in aplug-like manner downwardly through the vessel.

Further in accordance with the method that is more particularlydisclosed and described in U.S. Pat. No. 5,076,908 to Stangeland et al,a system for replacing catalyst during continuing operation of thenon-ebullating bed is assisted by carrying out the process at relativelyhigh liquid feed rates, even without ebullation of the bed. In apreferred form, the catalyst transfer system includes an inverted J-tubeas the withdrawal tube, so that the tube opens downwardly adjacent thecenter of the lower end of the vessel and directly above a centerportion of the surrounding annular flow paths of liquid and gas into thecatalyst bed. Thus, the intake for catalyst is out of the direct flow ofsuch streams, and particularly the gas flow. In such a preferred formthe annular flow paths are through a conical or pyramidal screen, orperforated plate, which supports the bed or column of catalyst acrossthe vessel through a plurality of radially spaced apart and axiallyelongated concentric rings, or polygons, supported by radial armsextending from the center of the vessel to the cylindrical side wall ofthe vessel. Each ring is formed by a pair of peripheral membersextending between the radial arms directly under the conical screen sothat this forms a circular gas pocket at the upper level in each ring sothat between each pair of said peripheral members alternate rings of gasand hydrocarbon liquid enter the bed simultaneously.

In accordance with a further preferred form of the method and apparatusthat is more particularly disclosed and described in Stangeland et al,catalyst is both withdrawn from the bed and added to the vessel underlaminar flow conditions as a liquid slurry to avoid abrasion and sizesegregation of particles during such transfer. Both the supply andwithdrawal flow lines have a minimum diameter of at least five timesand, preferably more than twenty times, the average diameter of thecatalyst particles. Further, the flow lines are of uniform diameterthroughout their length from either the catalyst supply chamber to thevessel, or from the vessel to the receiving chamber, including thethrough bore of a rotatable ball of the isolating, pressure controlvalves, known commercially as "full-port valves". Additionally, in eachcase a flush line is connected to the flow line between the isolatingvalve and the reactor vessel so that liquid hydrocarbon may be used toflush the line of catalyst or catalyst fines if necessary, before thevalve ball is closed. Preferably, but not necessarily, the withdrawalline may include means for flowing auxiliary hydrogen back into thereactor through the withdrawal tube to prevent coking due to hydrogenstarvation near or in the withdrawal tube.

The prior art does not disclose or suggest the above enumerated andpertinent features of either the total system or significant portions ofsuch a system in U.S. Pat. No. 5,076,908 to Stangeland et al, asdisclosed by the following patents:

Jacquin et al. U.S. Pat. No. 4,312,741, is directed toward a method ofon-stream catalyst replacement in a hydroprocessing system bycontrolling the feed of hydrogen gas at one or more levels. Catalyst, asan ebullated bed counterflows through the reactor but is slowed at eachof several levels by horizontally constricted areas which increase thehydrogen and hydrocarbon flow rates to sufficiently locally slowdownward flow of catalyst. While local recycling thus occurs at eachsuch stage, rapid through-flow of fresh catalyst, with resultant mixingwith deactivated or contaminated catalyst, is suppressed. The ebullatingbed aids simple gravity withdrawal of catalyst from the vessel.Improvement of the disclosed system over multiple vessels with valvesbetween stages is suggested to avoid the risk of rapid wear anddeterioration of valve seals by catalyst abrasion.

Kodera et al. U.S. Pat. No. 3,716,478, discloses low linear velocity ofa mixed feed of liquid and H₂ gas to avoid expansion (or contraction) ofcatalyst bed. By low linear velocity of fluid upflow, gas bubbles arecontrolled by flow through the packed bed, but the bed is fluidized byforming the bottom with a small cross-sectional area adjacent thewithdrawal tube. This assists discharge of catalyst without backmixingof contaminated catalyst with fresh catalyst at the top of the singlevessel. The range of the bed level in the vessel is from 0.9 to 1.1 ofthe allowable bed volume (±10%) due to fluid flow through the bed. Aparticular limitation of the system is that flow of the fluidsundergoing catalytic reaction is restricted to a rate that will notexceed such limits, but must be adequate to ebullate the bed adjacentthe catalyst withdrawal tube. Alternatively, injection of auxiliaryfluid from a slidable pipe section is required. The patenteesparticularly specify that the diameter of the lower end of the vessel issmaller to increase turbulence and ebullation of catalyst adjacent theinlet to the catalyst withdrawal line. Fluidization of catalyst isaccordingly indicated to be essential to the process. However thedisclosed gas flow rates are well below commercial flow rates and thereis no suggestion of temperatures or pressures used in the tests or thesize, density or shape of the catalyst.

Bischoff et al, U.S. Pat. No. 4,571,326, is directed to apparatus forwithdrawing catalyst through the center of a catalyst bed counterflowingto a liquid hydrocarbon and gas feed stream. The system is particularlydirected to arrangements for assuring uniform distribution of hydrogengas with the liquid feed across the cross-sectional area of the bed.Such uniform distribution appears to be created because the bed isebullating under the disclosed conditions of flow. Accordingly,considerable reactor space is used to initially mix the gas andhydrocarbon liquid feeds in the lower end of the vessel before flowingto other bottom feed distributors. The feeds are further mixed at ahigher level by such distributor means in the form of "Sulzer Plates" ora "honeycomb" of hexagonal tubes beneath a truncated, conical, orpyramidal-shaped funnel screen. The arrangement may include an open ramparea parallel to the underside of the screen between the tube or plateends. Further, to maintain gas distribution along the length of thecatalyst bed, quench gas is supplied through upflowing jets instar-shaped or annular headers extending across middle portions of thevessel. The arrangement for withdrawal of spent catalyst requiresebullation of at least the lower portion of the bed. As noted above,added vessel space for uniform mixing of hydrogen and feed beforeintroducing the fluids into an ebullated bed, as well as an ebullatingbed, increases the required size of the hydroprocessing vessel,increases catalyst attrition, increases catalyst bed mixing andsubstantially increases initial, and continuing operating costs of thesystem.

Bischoff et al. U.S. Pat. No. 4,639,354, more fully describes a methodof hydroprocessing, similar to U.S. Pat. No. 4,571,326, wherein similarapparatus obtains uniform ebullation through the vertical height of acatalyst bed, including a quench gas step.

Meaux U.S. Pat. No. 3,336,217, is particularly directed to a catalystwithdrawal method from an ebullating bed reactor. In the system,catalyst accumulating at the bottom of a vessel and supported on a flatbubble-tray may be withdrawn through an inverted J-tube having aparticular ratio of the volume of the short leg of the J-tube to thelonger leg. The diameter of the J-tube is suited only to flow ofcatalyst from a body of catalyst ebullated by the upflowing hydrocarbonfeed and gas.

U.S. Pat. Nos. 4,444,653 and 4,392,943, both to Euzen, et al., discloseremoval systems for catalyst replacement in an ebullating bed. In thesepatents, the fluid charge including hydrocarbon containing gas isintroduced by various arrangements of downwardly directed jets actinglaterally against or directly onto the conical upper surface of the bedsupport screen or screens. Alternatively, the feed is introduced througha conical screen after passing through a distributor arrangement oftortuous paths or a multiplicity of separate tubes to mix the gas andliquid feed over the conical screen. Such arrangements use aconsiderable volume of the pressure vessel to assure such mixing.

U.S. Pat. Nos. 3,730,880 and 3,880,569, both to Van der Toorn, et al.,disclose a series of catalytic reactors wherein catalyst movesdownwardly by gravity from vessel to vessel through check valves. Asnoted above, such valves require opening and closing to regulate therate of flow, or to start and stop catalyst transfer, with catalyst inthe valve flow path. Feed of process fluids is either co-current orcountercurrent through the catalyst bed.

Van ZijllLanghaut et al. U.S. Pat. No. 4,259,294, is directed to asystem for on-stream catalyst replacement by entrainment of the catalystin oil pumped as a slurry either to withdraw catalyst from or to supplyfresh catalyst to, a reactor vessel. Reacting feed is suggested to beeither co-current or countercurrent with catalyst flow through thereactor. Valves capable of closing with catalyst in the line, or afterback-flow of slurry oil, are required to seal off the catalystcontaining vessel at operating temperatures and pressures from thereceiving reactor vessel, or isolate the catalyst receiving lock hopperfrom the withdrawal section of the vessel.

Carson U.S. Pat. No. 3,470,900, and Sikama U.S. Pat. No. 4,167,474,respectively illustrate multiple single bed reactors and multi-bedreactors in which catalyst is replaced either continuously orperiodically. The feed and catalyst flow co-currently and/or radially.Catalyst is regenerated and returned to the reactor, or disposed of. Nocatalyst withdrawal system is disclosed apart from either theconfiguration of the internal bed support or the shape of the vesselbottom to assist gravity discharge of catalyst.

One of the basic principles and teachings of Stangeland et al in U.S.Pat. No. 5,076,908, is that by specifically selecting the size, shape,and density of the catalyst pellets, combined with appropriate controlof process liquid and gas velocities, random motion and backmixing ofthe catalyst can be minimized, and plugflow characteristics of thecatalyst downward and the liquid and gas flows upward, maximized.Stangeland et al economically utilizes space within a hydroprocessingvessel over a wide range of processing rates without substantial randommotion or ebullation of a packed bed of catalyst during high counterflowrates of the hydrocarbon feed and a hydrogen containing gas through thepacked bed, while maintaining continuous or intermittent replacement ofcatalyst for plug-like flow of the bed through the vessel. Such plugflow with high processing rates is obtained by Stangeland et al byselecting the size, shape and density of the catalyst particles toprevent ebullation and bed expansion at the design flow rate so as tomaximize the amount of catalyst in the vessel during normal operationand during catalyst transfer. Catalysts are selected utilizing datagained while studying catalyst bed expansion, such as in a large pilotplant run, with hydrocarbon, hydrogen and catalyst at the designpressures and flow velocities within the available reaction volume ofthe vessel. Catalyst is removed from the bed by Stangeland et al throughlaminar flow of the catalyst particles in a liquid slurry system inwhich the liquid flow line is uniform in diameter, and substantiallylarger than the catalyst particles, throughout the flow path between thereactor vessel and a pressurizable vessel including passageways throughthe flow control valves.

SUMMARY OF THE INVENTION

In accordance with one aspect of the present invention there is provideda method of periodically or semi-continuously transferring catalyst intoand out of a substantially packed bed of catalyst which is flowingdownwardly at a desired rate through a reactor vessel duringhydroprocessing over a wide range of counterflow rates of a hydrocarbonfeed stream that comprises a liquid hydrocarbon and ahydrogen-containing gas component which are flowing upwardly through thevessel. Such plug-like flow of the packed catalyst bed is achieved byselecting the average density, size, and shape of the catalyst particlesforming the catalyst bed so that the catalyst bed expands by less than10% by length at the maximum anticipated fluid flow velocities of thegaseous and liquid components therethrough. Desirably, such movement andbed level of such catalyst is continuously monitored to preventoverfilling and to assure minimum ebullation and attendant wastage ofreactor space and particle size segregation. Uniformity of gas flow ismaintained across the cross-sectional area of the vessel and the fullvolume of the bed so as to avoid ebullation of the bed, including eddycurrents or localized recirculation, of catalyst particles movingdownwardly in plug-like flow through the vessel. Preferably, the gaseouscomponent of the feed stream is uniformly distributed through aplurality of annular concentric rings, or polygons, formed by axiallyextending annular and radially spaced apart concentric supports under atruncated conical or pyramidal support screen. Such supports are axiallyelongated sufficiently to form a plurality of pairs of connected annulargas pockets and adjacent concentric liquid annular feed rings betweeneach adjacent pairs of annular supports. Thus, the catalyst bed isprovided with uniformly concentric annular and alternate feed rings ofboth liquid and gas across the full cross-sectional area of thedownwardly flowing catalyst bed.

In accordance with another aspect of the invention, the system forintroduction of quench gas at an intermediate level in the vesselmaintains the plug-like flow of catalyst downwardly through the vessel.Quench gas is introduced by a plurality of transversely extending pipemembers covered or shrouded by inverted V-shaped sheds. Each shedoverlies a quench gas supply pipe and acts to deflect catalyst outwardlyand downwardly over the apices. Each of the distributor sheds preferablyincludes a plurality of elongated slots along the lower edges of thetransverse sides. These slots form lateral gas redistribution channelsfor both upwardly flowing process gases and quench gas introducedthrough the transverse pipes.

Additionally the invention relates to methods and apparatus foron-stream replacement of catalyst without local levitation or ebullationof catalyst particles around the withdrawal point within the catalystbed by laminarly flowing a liquid hydrocarbon stream either into, or outof, the reactor vessel through a pair of flow paths. Each of the flowpaths has a substantially constant cross-sectional area throughout itslength and a diameter at least five times the average diameter of thecatalyst particles flowing between said vessel and at least one andpreferably two, pressurizable catalyst lock-hoppers or receptacles,serving respectively to supply fresh catalyst to the top of the bed andto remove spent catalyst from the bottom. Further, each flow pathincludes at least one in-line control valve having a through-bore ofsubstantially the same diameter as the flow path and at least oneauxiliary fluid flow path for introducing fluid flow into the slurrypath for flushing catalyst particles from the path. Preferably, theflush fluid is a liquid, and selectively, by reverse hydrogen flowthrough the line when catalyst is not being moved, particles are blockedfrom entering the flow path and coking is prevented at the entry to theflow tube. The catalyst vessels are selectively pressurizable asrequired to induce such laminar flow of liquid-entrained catalyst tofeed replacement catalyst into the upper end of the reactor vessel andto withdraw spent catalyst from the lower end of the vessel. Desirably,each of the flow paths is characterized by an inverted J-tube, whichincludes an inlet portion for the liquid stream and entrained catalysthaving a reverse upward flow section substantially shorter than thedownward flow path. Preferably, in the reactor vessel the inlet portionfor withdrawing catalyst is disposed above an unperforated centralportion of the conical bed support screen so that such catalystwithdrawal position is adjacent the bed bottom, but substantially out ofa plurality of concentric feed paths for upwardly flowing liquidhydrocarbon feed and gas streams. This avoids gas entrainment with thecatalyst slurry, as by ebullation of the bed around the intake point.

In a preferred embodiment of the invention, the present inventionaccomplishes its desired objects by broadly providing a catalystcomprising a plurality of catalytic particulates having a mean diameterranging from about 35 Tyler mesh to about 3 Tyler mesh; and a sizedistribution such that at least about 90% by weight of the catalyticparticulates have a diameter ranging from R₁ to about R₂, wherein:

(1) R₁ has a value ranging from about 1/64 inch to about 1/4 inch,

(2) R₂ has a value ranging from about 1/64 inch to about 1/4 inch, and

(3) a value of a ratio R₂ /R₁ ranges from about 1.0 to about 1.4, and anaspect ratio of less than about 2.0. The catalyst may be employed in anyhydrogenation process. Preferably, the catalyst is for producing aplug-flowing substantially packed bed of hydroprocessing catalyst duringhydroprocessing by contacting a substantially packed bed ofhydroprocessing catalyst with an upflowing hydrocarbon feed stream. Moreparticularly, when the catalytic particulates are disposed in ahydrocarbon reaction zone, a substantially packed bed of hydroprocessingcatalyst is produced; and when a hydrocarbon feed stream flows upwardlythrough the substantially packed bed of hydroprocessing catalyst,plug-flowing commences when a volume of the catalytic particulates iswithdrawn from a bottom of the hydrocarbon reaction zone. As used herein"catalyst" includes other particles which interact with a feed stream,such as sorbents, or other fluid contact bodies. The catalyst isdisposed in a reaction zone and a hydrocarbon feed stream is flowedupwardly through the catalyst for hydroprocessing the hydrocarbon feedstream.

The present invention also accomplishes its desired objects by broadlyproviding a method for hydroprocessing a hydrocarbon feed stream that isupflowing through a hydroconversion reaction zone having a substantiallypacked bed of catalyst comprising the steps of:

(a) forming a plurality of annular mixture zones under a hydroconversionreaction zone having a substantially packed bed of hydroprocessingcatalyst such that each of the annular mixture zones contains ahydrocarbon feed stream having a liquid component and ahydrogen-containing gas component and wherein the annular mixture zonesare concentric with respect to each other and are coaxial with respectto the hydroconversion reaction zone; and

(b) introducing the hydrocarbon feed stream from each of the annularmixture zones of step (a) into the substantially packed bed ofhydroprocessing catalyst to commence upflowing of the hydrocarbon feedstream from each of the annular mixture zones through the substantiallypacked bed of the catalyst.

The present invention also accomplishes its desired objects by broadlyproviding a method for hydroprocessing a hydrocarbon feed stream that isupflowing through a hydroconversion reaction zone having a substantiallypacked bed of catalyst comprising the steps of:

(a) disposing catalyst in a reaction zone, said catalyst comprising aplurality of catalytic particulates having a mean diameter ranging fromabout 35 Tyler mesh to about 3 Tyler mesh and a size distribution suchthat at least about 90% by weight of said catalytic particulates have adiameter ranging from R₁ to R₂, wherein:

(1) R₁ has a value ranging from about 1/64 inch to about 1/4 inch,

(2) R₂ has a value ranging from about 1/64 inch to about 1/4 inch,

(3) a value of a ratio R₂ /R₁ ranges from about 1.0 to about 1.4, and anaspect ratio of less than about 2.0; and

(b) upflowing through the catalyst of step (a) a hydrocarbon feed streamfor hydroprocessing the hydrocarbon feed stream.

The catalytic particulates have a size distribution such that a maximumof about 2.0% by weight of said catalytic particulates have a diameterless than R₁. The catalytic particulates also have a size distributionsuch that a maximum of about 0.4% by weight of the catalyticparticulates have a diameter less than R₃, wherein R₃ is less than R₁and the value of the ratio R₁ /R₃ is about 1.4. The catalyticparticulates have a maximum attrition of about 1.0% by weight of thecatalytic particulates through a diameter having a value of R₁ ; and thecatalytic particulates have a maximum attrition of about 0.4% by weightof the catalytic particulates through a diameter having a value of R₃,wherein R₃ is less than R₁ and the value of the ratio R₁ /R₃ is about1.4.

In another aspect of the invention, the present invention alsoaccomplishes its desired objects by broadly providing a method forproducing an essentially downwardly plug-flowing substantially packedbed of hydroprocessing catalyst within a hydroconversion reaction zonecomprising the steps of:

(a) forming a plurality of annular mixture zones under a hydroconversionreaction zone having a substantially packed bed of hydroprocessingcatalyst such that each of the annular mixture zones contains ahydrocarbon feed stream having a liquid component and ahydrogen-containing gas component and wherein the annular mixture zonesare concentric with respect to each other and are coaxial with respectto the hydroconversion reaction zone;

(b) introducing the hydrocarbon feed stream from each of the annularmixture zones of step (a) into the substantially packed bed ofhydroprocessing catalyst to commence upflowing of the hydrocarbon feedstream from each of the annular mixture zones through the substantiallypacked bed of the catalyst;

(c) withdrawing a volume of particulate catalyst from thehydroconversion reaction zone to produce an essentially downwardlyplug-flowing substantially packed bed of hydroprocessing catalyst withinthe hydroconversion reaction zone.

In another broader aspect of the invention, the present invention alsoaccomplishes its desired objects by broadly providing a method formaximally occupying a reactor volume with a substantially packed bed ofhydroprocessing catalyst during hydroprocessing by contacting thesubstantially packed bed of hydroprocessing catalyst with an upflowinghydrocarbon feed stream having a liquid component and ahydrogen-containing gas component. The method for maximally occupying areactor volume with a substantially packed bed of hydroprocessingcatalyst preferably comprises the steps of:

(a) disposing a substantially packed bed of hydroprocessing catalyst ina reactor zone (or reaction zone or zone for reaction) contained withina reactor volume;

(b) upflowing into the substantially packed bed of hydroprocessingcatalyst a hydroprocessing feed stream including a liquid component anda hydrogen-containing gas component and having a rate of flow such thatthe expansion of the substantially packed bed of hydroprocessingcatalyst is limited to less than 10% by length beyond a substantiallyfull axial length of the substantially packed bed of hydroprocessingcatalyst in a packed bed state;

(c) withdrawing a volume of the hydroprocessing catalyst from thereactor zone to commence essentially plug-flowing downwardly of thesubstantially packed bed of hydroprocessing catalyst within the reactorzone; and

(d) adding hydroprocessing replacement catalyst to the essentiallyplug-flowing downwardly, substantially packed bed of hydroprocessingcatalyst at a volume and/or rate to substantially replace the volume ofthe withdrawn hydroprocessing catalyst. The substantially packed bed ofhydroprocessing catalyst is disposed in the reactor zone within thereactor volume such that the substantially packed bed of hydroprocessingcatalyst maximally occupies the reactor volume. The substantially packedbed of hydroprocessing catalyst occupies at least about 50% by volume ofthe reactor volume; preferably at least about 60% by volume; and morepreferably at least about 65% or 70% by volume of the reactor volume.

From the foregoing summary it will be apparent that several significantfactors contribute directly to the present invention accomplishing itsdesired objects, and to the efficient use of a given process reactorvessel to assure non-ebullating, plug-like flow of a body of catalystparticles therethrough while being contacted by a counter-flowinghydrocarbon fluid stream of gas and liquid at maximum space-velocity.Among such significant factors are: (i) the size, volume and densitycharacteristics of such catalyst particles at preselectable flowvelocities and pressures of the hydrocarbon fluid stream; (ii) controlof catalyst bed ebullation and/or levitation during hydrocarbon fluidand hydrogen flows; (iii) laminar flow of the catalyst particles duringmovement into and out of the catalyst moving bed for replacement (orregeneration or rejuvenation) without bed ebullation or levitation; (iv)concentric annular feed of alternate rings of the gas and liquidcomponents of the hydrocarbon feed uniformly into the full movingcatalyst bed, which is capable of recovering promptly from upset orpressure changes in the reactor vessel to restore such alternate ringsof gas and liquid over process runs of extended length (e.g. severalthousand hours); and (v) redistribution of the gas components along theaxial length of the moving bed.

It is another object of the present invention to broadly provide amethod for producing an essentially downwardly plug-flowingsubstantially packed bed of hydroprocessing catalyst within ahydroconversion reaction zone.

These, together with the various ancillary objects and features whichwill become apparent to those skilled in the art as the followingdescription proceeds, are attained by this invention, a preferredembodiment as shown with reference to the accompanying drawings, by wayof example only, wherein:

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 is a schematic view of a typical hydroprocessing vessel to whichthe present invention is particularly directed for on-stream catalystreplacement during continuous plug-like flow of catalyst throughup-flowing liquid hydrocarbon feed and gas streams;

FIG. 2 is a bottom plan view of the concentric and radial catalyst bedsupport means for a truncated conical or pyramidal screen, taken in thedirection of arrows and along the plane of line 2--2 in FIG. 1.;

FIG. 3 is an elevational cross-sectional view of the support means andscreen taken in the direction of arrows and along the plane of line 3--3in FIG. 2;

FIG. 4 is a partial elevational view of an alternate form of a laminarflow arrangement for withdrawing deactivated catalyst particles from thereactor bed;

FIG. 5 is a cross-sectional plan view of the reactor vessel taken in thedirection of arrows and along the plane of line 5--5 in FIG. 1 showing apreferred form of gas redistribution and quench system over a centralportion of the catalyst bed;

FIG. 6 is a perspective view, partially in cross-section, of one of thequench or redistribution shed units shown in FIG. 5;

FIG. 7 is a perspective view of a preferred arrangement of two tiers ofShed units of FIG. 5 at a given level in the catalyst bed;

FIG. 8 is a partial cross-sectional view illustrating a catalytic bedwith a plurality of superimposed layers with respect to each otherbefore commencement of a plug-flow;

FIG. 9 is a partial cross-sectional view illustrating a catalytic bedwhich is moving downwardly in a plug-flow fashion;

FIG. 10 is a partial cross-sectional view of the reactor and a partialperspective view of another embodiment of the catalytic support means;

FIG. 11 is a partial cross-sectional view of the reactor and thecatalytic support means of FIG. 10 which includes a plurality of annularmixture zones under the substantially packed bed of hydroprocessingcatalyst with each annular mixture zone containing a liquid hydrocarboncomponent and a hydrogen-containing gas component and wherein theannular mixture zones are concentric with respect to each other and arecoaxial with respect to the reactor and the substantially packed bed ofhydroprocessing catalyst;

FIG. 12 is the partial cross-sectional view of the reactor and supportmeans in FIG. 11 with the inert pellets, and illustrating ribs or spokessecured to an imperforate center plate and supporting a plurality ofsegmented plates; and

FIG. 13 is another cross-sectional view of the reactor and support meansas similarly illustrated in FIG. 12 with a bed of inert pellets having aliquid hydrocarbon component and a hydrogen-containing gas componentflowing around the inert pellets for entering the annular mixture zones.

DETAILED DESCRIPTION OF THE INVENTION INCLUDING PREFERRED EMBODIMENTS OFTHE INVENTION

Referring in detail now to the drawings, and initially more particularlyto FIG. 1, a hydroprocessing system is shown embodying the method of thepresent invention to increase substantially both the continued catalyticactivity of a volume or bed of catalyst 10 and the efficient use of asingle reactor vessel of a given reactor volume, such as reactor vessel11. Vessel 11, as indicated by the thickness of its cylindrical sidewall 12 and domed closure heads, or ends, 13 and 14, is designed toreact a hydrogen containing gas mixed with a liquid hydrocarbon streamat a pressure of up to about 300 atmospheres (about 4500 lbs per squareinch) and up to about 650° C. (about 1200° F.). Such reaction gas and afeed stream of hydrocarbon liquids are preferably premixed andintroduced as a single stream through bottom head 13 by line 16.

To assure maximum catalytic benefit during the hydroprocessing of thehydrocarbon feed stream and the hydrogen-containing gas, it is essentialthat vessel 11 contain as much catalyst as possible within the designvolume of vessel 11. Accordingly as indicated, support means 17 forcatalyst bed 10 is placed as low as possible in vessel 11 while assuringfull and adequate dispersion of the hydrogen phase within the liquidhydrocarbon stream. At the same time, the upper limit of bed 10 is nearthe top of domed head 14, while providing an adequate space 21 fordisengaging any entrained catalyst from the resulting products withdrawnthrough center pipe be. To insure that catalyst is not entrained intoproduct fluids exiting through center pipe 18, a screen 15 may beinstalled in space 21 above a bed surface 20 defining the top of thecatalyst bed 10. Fresh catalyst is then added to bed surface 20 throughpipe 19 extending through screen 15. Desirably, the upper level or topof the catalyst bed 10, designated as the bed surface 20, is preferablycontrolled on a continuous basis by gamma ray absorption measurementmade possible by a gamma ray source 22 and gamma ray detector 24positioned in close proximity to the bed surface 20 of catalyst bed 10.Such a gamma ray source may be in the form of radioactive isotopes, suchas Cesium 137, disposed inside the reactor in a specially designed well.Alternatively the source can be an electrically controllable source,such as a thermal neutron activated gamma ray generator. Detector 24 maybe in the form of an ionization tube, Geiger-Mueller tube or ascintillation detector. Suitable sources and detectors are manufacturedby Ronan Engineering Co., Texas Nuclear and other vendors. By detectingthe level of surface 20, it is possible, in accordance with theinvention, to insure that the catalyst inventory is maintained at theoptimum level and that the reactor is never overfilled.

Overfilling the reactor increases the chance that catalyst particleswill be crushed in the isolation valves in the transfer lines when theyare closed, at the end of each transfer. Bed level control is alsoneeded to confirm that ebullation of the bed is minimized and thatundesirable excursions from the design flow rate for hydrogen andhydrocarbon feed flowing upwardly through bed 10 are avoided for theselected catalyst. To this end, the size, shape, and density of catalystparticles supplied to the bed are selected in accordance with thedesigned maximum rate of flow of the feed streams to prevent suchebullation. Such control assures that bed 10 progressively moves downthrough vessel 11 in layers as by a plug flow. A "plug flow" of thecatalyst bed 10 is illustrated in FIGS. 8 and 9 and may be bestdescribed as when a lowermost volumetric layer A is removed, the nextvolumetric layer B flows downwardly to replace the lowermost volumetriclayer B and assumes a new position as a lowermost volumetric layer B.The removed lowermost volumetric layer A is replaced with an uppervolumetric layer J. The procedure is again repeated (as best shown bythe dotted line representations in FIG. 9) by removing the lowermostvolumetric layer B and causing the next volumetric layer C to flowdownwardly in a plug-like fashion to replace the lowermost volumetriclayer B and assume a new position as a lowermost volumetric layer C. Theremoved lowermost volumetric layer B is replaced with an uppervolumetric layer K. The procedure may be continually repeated to definea downwardly plug-flowing catalyst bed 10 which is moving in directionof arrow W in FIG. 9.

The procedure to determine whether or not a catalyst bed 10 isplug-flowing may be by any suitable procedure. For example, in apreferred embodiment of the present invention wherein metals (e.g.vanadium) are being removed from a hydrocarbon feed stream, the catalystbed 10 is plug-flowing if a catalytic sample (e.g. 15 catalyticparticulates) from withdrawn catalyst is analyzed and it is foundthrough elemental metal analysis that the catalytic sample has a uniformhigh metal load, preferably at least about 1.5 times more than theaverage metal load of the catalyst bed 10, and more preferably at leastabout 2.0 times more than the average metal load of the catalyst bed 10.Those possessing the ordinary skill in the art can determine the averageload of the catalyst bed 10 from the total amount of metals removed fromthe hydrocarbon feed stream, the weight of the catalytic bed 10, etc.

It is to be understood that whenever the specification or the claimsstates or mentions any type of catalyst movement or catalyst bed 10movement (e.g. "removing", "moving", "supplying", "replacing","delivering", "flow", "flowing", "transfer", "transferring", "addition","adding", "admixing", etc.) for any type or mixture of catalyst withoutstating or mentioning the basis, the basis for such type of catalyst orcatalyst bed movement may be on any type of basis, such as "intermittentbasis", "periodic basis", "continuous basis", "semi-continuous basis",etc. Thus, by way of example only, removal of lowermost volumetriccatalytic layers and addition of upper volumetric catalytic layers maybe on a "periodic basis", "a continuous basis", or even "a one timebasis", all without affecting the spirit and scope of the presentinvention(s). It is to be also understood that the "removal" or"withdrawal" of catalyst and the "addition" or "replacement" of catalystare mutually exclusive of each other and may be performed simultaneouslyor at different times without affecting the spirit and scope and of thepresent invention(s). Preferably, the "addition" or "replacement" ofcatalyst is performed after the "removal" or "withdrawal" of catalystand after the catalyst bed 10 has moved downwardly into a non-movingstate or non-moving posture.

Catalysts are selected utilizing data acquired by measuring bedexpansion, such as in a large pilot plant run, with hydrocarbon,hydrogen and catalyst as described below and illustrated in Example 2.

To further assure that plug flow continues throughout the full length ofthe bed, and particularly at the bottom portion, bed support means 17 isparticularly characterized by the truncated polygonal or conicalconfiguration of support means 17.

As shown in the preferred embodiment of FIGS. 2 and 3, and as best seenin FIG. 2, support 17 includes a series of annular polygons, approachingthe form of annular rings, formed by a plurality of segment plates 27between radial ribs or spokes 26 extending from imperforate center plate25 to sidewall 12 of vessel 11. As shown in FIG. 3, spokes 26 may be anysuitable geometric shape, such as rod-like (see FIG. 3) or substantiallyflat plates (see FIG. 10), which divide the circumference of the vesselinto many segments (eight in this case) and similarly support the endsof outer octagonal ring 23 of support means 17 formed by annular orcircumferential plates 27. In each case, radial ribs or spokes 26, andannular segment plates 27 form a plurality of concentric rings, orannular polygons which support conical, or pyramidal, perforated plateor screen 28. Thus screen 28 is permeable to both gas and liquid risingfrom the lower portion of vessel 11.

In one preferred embodiment of the particular merit of the concentricannular polygons as illustrated in FIG. 3, the interconnected plates 27form a plurality of ring-like structures extending generally axiallyparallel to the sidewall 12 with the radial ribs or spokes 26 radiallyextending towards the sidewall 12 of reactor vessel 11. The mixture ofthe hydrocarbon liquid feed and hydrogen gas that is to enter thecatalyst bed 10 separates by gravity into radially alternate gas andliquid rings, made up of adjacent segments between each pair or radialspokes 26. Thus, both phases flow upwardly through alternate concentricannular passages under screen 28. The preferential separation of gasfrom liquid in each ring includes an annular cap segment of gasoverlying an adjacent lower annular segment filled with liquid. Hence,both fluids have equal, and angularly adjacent, access to the bedthrough screen 28. The plurality of alternate annular rings of hydrogengas and hydrocarbon liquid assure even and equal feed of both phasesacross the full cross-sectional area of screen 28 into bed 10. Amongother factors, we have particularly found that this configurationinsures even and equal distribution across the full cross- sectionalarea of the catalyst bed. Such equal distribution across the fulldiameter of the bed 10, permits a quiescent flow section to formdirectly above center plate 25 which truncates conical bed support means17. This decreases substantially potential local ebullation or eddycurrents from being induced in the catalyst bed at the point of catalystwithdrawal through inlet 30 of inverted J-tube 29 to assure localizedlaminar flow of catalyst and liquid from within bed 10.

Uniform feed of the mixture of the hydrocarbon feed stream and hydrogenis particularly facilitated to the inlet side of plates 27 of supportmeans 17 through plenum or inlet chamber 33 enclosed between support 17and circular plate member 31, which extends across the fullcross-sectional area of vessel 11. The circular plate member 31 definesa grid-like structure for supporting a permeable screen 6 having one ormore openings, as best shown in FIGS. 11, 12 and 13. As further bestshown in FIGS. 11, 12 and 13, the permeable screen 6 supports a bed 3 ofa plurality of inert pellets 4 (e.g. alumina pellets) which are sizednot to pass through the openings in the permeable screen 6, to preventeddy currents in the plenum chamber 33, and to keep bubbles ofhydrogen-containing gas diffused within the hydrocarbon feed streams.Plate 31 includes a multiplicity of similar large diameter tubes 32forming openings through plate 31. Each tube is several inches indiameter and extends axially to a similar depth, say on the order of 4to 6 inches, below plate 31. Tubes 32 provide equal access to themixture of hydrogen and hydrocarbon feed stream into plenum chamber 33.Even distribution of the incoming feed stream into bottom header 35 fromfeed line 16 may also be assisted by deflector plate 34 to assure thatoversized bubbles of hydrogen that may be contained in the feed streamwill be equally distributed across the full cross-sectional area ofplate 31 and equally distributed to each of tubes 32 for flow intoplenum chamber 33. The length of tubes 32 may be selected to form asuitable gas head under plate 31 to suppress surges in the feed streamsentering header 35.

As noted above, the vertical, transverse width or axial length of plates27 which set off each individual annular and radial segment, provideequal access to both hydrogen and liquid feed into catalyst bed 10, andstepped under screen 28 so that they effectively form rings of gas andhydrocarbon feed alternately across the full diameter at the inlet sideof catalyst bed 10. In this way, no single area of the inlet to catalystbed 10 becomes a segregated or preferential, flow path for either gas orthe liquid. Further, if pressure surges result in full wetting of screen28 by the liquid phase, recovery of gas flow is assisted by the arealbreadth of each segment between plates 27 and radial plates 26.

In another preferred embodiment of the particular merit of theconcentric annular polygons as illustrated in FIGS. 10, 11, 12 and 13,there is seen a liquid hydrocarbon component LH and ahydrogen-containing gas component HG (hydrogen-containing gas bubbles)entering as an LH-HG mixture into the plenum chamber 33 from tubes 32.The LH-HG mixture is introduced into the plenum chamber 33. In thispreferred embodiment of the present invention, the annular orcircumferential plates 27 are secured to and are supported by the radialribs or spokes 26, each of which has a vertical or transverse width thatis essentially equal to the vertical or transverse width of the annularor circumferential plates 27. The radial ribs or spokes 26 also functionas a means for reducing a size of hydrogen-containing gas bubbles,especially over-size hydrogen-containing gas bubbles from thehydrogen-containing gas component HG. Those skilled in the art willreadily recognize that the number of radial ribs or spokes 26 employedwill depend on a number of factors, such as the anticipated number ofover-size hydrogen-containing gas bubbles in the upwardly flowinghydrocarbon feed stream, the weight of the catalyst bed 10, etc. Theinterconnected plates 27 and radial ribs or spokes 26 form a web orweb-like structure defining a plurality of annular mixture zones,generally illustrated as MZ in FIGS. 10 and 11. The annular mixturezones MZ are essentially continuous or are generally endless annularmixture zones MZ, and may contain or be subdivided into any reasonabledesired number of mixture zones (or sub-mixture zones), such as MZ₁,MZ₂, MZ₃, MZ₄, MZ₅, and MZ₆ in FIGS. 10 and 11. Each of the individualmixture zones MZ₁, MZ₂, MZ₃, MZ₄, MZ₅, and MZ₆ is for all practicalpurposes an annularly continuous or endless mixture zone of uniformthickness, excepting a periodic interruption by radially ribs 26, whichare relatively narrow vis-a-vis the spaced distance between any pair ofcontiguous ribs 26--26. As evident in FIGS. 10, 11, 12 and 13,concentric with mixture zone MZ₁ and as a partial bottom to same isimperforate center plate 25, which is preferably spaced from and off ofthe plate 31 and the screen 6 such that inert pellets 4 may be supportedby the screen 6 and the plate 31 immediately underneath the imperforatecenter plate 25. Mixture zone MZ₁ is essentially a cylindrical annularmixture zone with an open top and boundaries defined by the spacebetween a plurality of interengaged and coupled plates 27₁ s and theperimeter of the imperforate center plate 25.

The plurality of annular mixture zones MZ (or the annularly continuousor endless mixture zones MZ₂ s, MZ₃ s, MZ₄ s, MZ₅ s, and MZ₆ s) underthe catalyst bed 10 are concentric with respect to each other and arecoaxial with respect to the reactor vessel 11 and the catalyst bed 10.The plates 27 may be radially spaced from each other at any suitabledistance (preferably of uniform distance) to assist in accomplishing thedesired objects of the present invention; however, preferably the plates27 are radially spaced from each other at a generally uniform thicknessor distance that ranges from about 1 inch to about 4 feet, morepreferably from about 6 inches to about 3 feet, most preferably fromabout 1 foot to about 2 feet. The radially spaced relationship betweenand among the plates 27 generally defines a uniform thickness for eachof the mixture zones (i.e. MZ₂ s, MZ₃ s, etc.). It is to be understoodthat while the plurality of annular mixture zones MZ is represented inFIGS. 2, 3, 10, 11, 12 and 13 as being a plurality of non-circulargeometric-shaped zones (e.g. octagonal in FIGS. 2), the spirit and scopeof the present invention includes that the plurality mixture zones MZmay comprise any geometric-shaped zones including not onlypolygonal-shaped zones, but also a plurality of concentric circularmixture zones, etc., all of which would also be concentric with respectto each other and coaxial with respect to the reactor vessel 11 and/orthe catalyst bed 10 (or the hydroconversion reaction zone).

Therefore, the plates 27 function to form generally uniform thick andessentially circular bands of concentric hydrocarbon feed streams thatare also coaxial with respect to the catalyst bed 10. By way of exampleonly and as best shown in FIGS. 2 and 10, angular mixture zone MZ₂ isdefined by the eight (8) interengaged or intercoupled plates 27₁ s andthe eight (8) interengaged or intercoupled plates 27₂ s. The eight (8)plates 27₁ s and the eight (8) plates 27₂ s each form an annulateboundary for the essentially circular band of hydrocarbon feed stream inmixture zone MZ₂. Because the spacing or distance between plates 27₁ sand 27₂ s is generally circumferentially uniform, the thickness or sizeof the essentially circular band of hydrocarbon feed stream in mixturezone MZ₂ is essentially uniform transversely and/or equal in transverseor horizontal cross section. Similarly, mixture zone MZ₆ is defined bythe eight (8) interengaged or intercoupled plates 27₅ s and the eight(8) interengaged or intercoupled plates 27₆ s, the combination of whichform annulate boundaries for the essentially circular band ofhydrocarbon feed stream in mixture zone MZ₆. As was previously similarlyindicated for plates 27₁ s and 27₂ s, because the spacing or distancebetween plates 27₅ s and 27₆ s is generally circumferentially uniform,the thickness or size of the circular band of hydrocarbon feed stream inmixture zone MZ₆ is essentially uniform transversely and/or equal intransverse or horizontal cross section. Plates 27₂, 27₃, 27₄, and 27₅similarly functionally interengage and intercouple to define annulateboundaries for mixture zones MZ₃, MZ₄, and MZ₅. As indicated and as bestshown in FIG. 2, ribs 26 extend radially from imperforate center plate25 and planarly represent visually pie-shaped segments. Between any pairof contiguous ribs 26--26, the lengths of the respective plates 27increase from plate 27₁ to or towards plate 27₆ while the widths areessentially the same as best shown in FIG. 3. Thus, plate 272 is longerthan plate 27₁ while possessing the identical approximate width.Likewise: plate 27₃ is longer than plate 27₂, plate 27₄ is longer than27₃, plate 27₅ is longer than plate 27₄, and plate 27₆ is longer thanplate 27₅, while all the plates 27 simultaneously have generally thesame width or the same longitudinal extension below the screen 28 (seeFIG. 3). Thus, the vertical dimensions or the widths of the plates 27(i.e. the structural extensions of the plates 27 that are generallyparallel to the longitudinal axis of the reactor vessel 11 and/or thecatalyst bed 10 therein) are generally equal. All plates 27 arepreferably spaced such that the hydrocarbon feed stream flows parallelto the longitudinal axis of the catalyst bed 10 before contacting andentering the same. Both the upper edges and lower edges of plates 27₁ s,27₂ s, 27₃ s, 27₄ s, 27₅ s, and 27₆ s are all at a different level orheight, as best shown in FIGS. 3 and 11. The mixture zones MZ differfrom a plurality of tubes, conduits, or pipe-like passages forintroducing an essentially complete or essentially total integralcylindrical hydrocarbon feed stream into the catalytic bed 10. As bestshown in FIGS. 3 and 11, the upper and lower edges of plates 27₁ s areat a different level or height than the upper and lower edges of plates27₂ s which are at a different level or height than the upper and loweredges of plates 27₃ s. Similarly, the upper and lower edges of plates27_(3s) are at a different level or height than the upper and loweredges of plates 27₄ s which are at a different level or height than theupper and lower edges of plates 27₅ s. The upper and lower edges of thelatter are at a different level or height than the upper and lower edgesof plates 27₆ s.

After the LH-HG mixture enters and flows through the screen 6 into theplenum chamber 33, the flowing LH-HG mixture enters into each of thegenerally continuous annular mixture zones MZ₂ s, MZ₃ s, etc. fordividing or separating the flowing LH-HG mixture into a plurality offlowing generally continuous annular LH-HG mixtures, which have beendesignated LH-HG₂, LH-HG₃, LH-HG₄, LH-HG₅ and LH-HG₆ in FIG. 11. As waspreviously indicated, the mixture zone MZ₁ is also basically an annularor cylindrical shaped mixture zone defined by the space between theperimeter of the imperforate center plate 25 and intercoupled segmentedplates 27₁ s and receives hydrocarbon feed stream (i.e. hydrocarbonliquid feed and/or hydrogen gas) in and through the space by which it isbeing defined. In a preferred embodiment of the present invention and asbest shown in FIG. 13, before the flowing LH-HG mixture enters into eachof the generally continuous annular mixture zones MZ₁ s, MZ₂ s, MZ₃ s,etc. the LH-HG mixture flows around the plurality of inert pellets 4 inzig-zag fashions for reducing the possibility of eddy currents and forkeeping bubbles of hydrogen gas diffused within the liquid hydrocarbonand preventing agglomeration of same into larger size bubbles. Thehydrocarbon feed stream entering into mixture zone MZ₁ is designatedLH-HG₁. The plurality of LH-HG mixtures (i.e. LH-HG₁, LH-HG₂, etc.) passthrough the screen 28 and respectively enter into the catalyst bed 10from each of the mixture zones (i.e. MZ₁ s, MZ₂ s, MZ₃ s, etc.) at aflow rate such as not to ebullate, levitate or expand the catalyst bed10 upwardly and/or towards the screen 15 and the domed head 14 by morethan 10% by length beyond substantially the full axial length of the bedcatalyst 10 in a packed bed state, such as the packed bed statereflected in FIG. 8. The plurality of generally continuous annular LH-HGmixtures flow upwardly through screen 28 and into the catalyst bed 10.The catalyst bed 10 in the present invention preferably comprisescatalyst particles which are substantially the same and/or uniform size,shape, and density and which are selected in accordance with the averageoptimum velocity of the hydrocarbon feed stream (i.e. a mixture of aliquid hydrocarbon component LH and a hydrogen-containing gas componentHG, or the continuous annular LH-HG mixtures) flowing into the plenumchamber 33 and subsequently into and through the plurality of mixturezones MZ₂ s, MZ₃ s, etc. The rates of flow of the plurality of therespective LH-HG mixtures (i.e. LH-HG₁, LH-HG₂, etc.) from therespective mixture zones MZ₁ s, MZ₂ s, etc., and thus also the flowrates of the hydrocarbon feed stream into plenum chamber 33 from andthrough line 16, are all to be controlled in an amount and to an extentsufficient to maintain expansion or levitation of the catalyst bed 10 toless than 10% by length over or beyond substantially the full axiallength of the bed 10 in a packed bed state. More preferably, theexpansion of the substantially packed bed of catalyst is limited to lessthan 5%, most preferably less than 2% or even less than 1%, by lengthover or beyond substantially the full axial length of the bed 10 in apacked bed state. Ideally the expansion of the substantially packed bedof catalyst is limited to essentially 0% by length.

The flow rate of the hydrocarbon feed stream through line 16 is to be ata rate not substantially greater than the optimum rate of flow. Theoptimum rate of process fluid flow through the substantially packed bedof catalyst will vary from process unit to process unit based on severalfactors including oil and hydrogen feed characteristics, catalystspecifications, process objectives, etc. Based on catalyst particleshaving substantially the same and/or uniform size, shape and density,the flow rate of the hydrocarbon feed stream preferably ranges fromabout 0.01 ft/sec to about 10.00 ft/sec and more preferably from about0.01 ft/sec to about 1.00 ft/sec. Similarly and/or likewise and furtherbased on the catalyst particles having substantially the same and/oruniform size, shape, and density, the flow rate of the continuousannular LH-HG mixtures (i.e. the summation of the flow rates for LH-HG₁through LH-HG₆ from mixture zones MZ₁ s through MZ₂ s respectively inFIG. 11) is to be at a rate not substantially greater than the optimumrate of flow, preferably ranging from about 0.01 ft/sec to about 10.00ft/sec, and more preferably from about 0.01 ft/sec to about 1.00 ft/sec.The specific flow rate would depend as indicated on a number ofvariables, such as the particular application (e.g. demetallation ordesulfurization etc.) of the hydroprocessing process. The specific flowrates however would be at any suitable rate controlled in an amount andto an extent sufficient to limit expansion of the substantially packedbed of catalyst to less than 10% by length over or beyond asubstantially packed bed of hydroprocessing catalyst in a packed bedstate.

In a preferred embodiment of the invention and for such a flow rate forthe hydrocarbon feed stream and for such a flow rate for the continuousannular LH-HG mixtures, the catalyst particles preferably have thesubstantially same and/or uniform size, shape and density in order toobtain over the desired demetalization and/or desulfurization of theliquid hydrocarbon component LH in the hydrocarbon feed stream (i.e.LH-HC mixture) into produced hydrogen upgraded product fluids that arebeing withdrawn from the reactor vessel 11 through the center pipe 18.At the above indicated flow rates for the hydrocarbon feed streamflowing through line 16, and for the flow rates for the generallycontinuous annular LH-HG mixtures (i.e. LH-HG₁, LH-HG₂, etc.), theproduced upgraded product fluids are being preferably withdrawn throughthe center pipe be from the reactor vessel 11 at a rate ranging fromabout 0.01 ft/sec to about 10.00 ft/sec and more preferably from about0.01 ft/sec to about 1.00 ft/sec. The withdrawal rate(s) of the producedupgraded product fluids is not to be greater than the optimum rate offlow and will also vary from process unit to process unit based onseveral factors including oil and hydrogen feed characteristics,catalyst specifications, process objectives, etc. The specificwithdrawal rate(s) would be any suitable withdrawal rate, controlled inan amount and to an extent sufficient to prevent and/or limit expansionof the substantially packed bed of catalyst to less than 10% (morepreferably less than 5%, most preferably less than 2% or even less than1%) by length over or beyond substantially the full axial length of thebed 10 in a packed bed state.

The arrangement in inlet distributor 31 for uniformly distributinghydrogen gas and liquid hydrocarbon feed as shown in FIG. 3 may bemodified by lengthening or shortening tubes 32, forming uniformlydistributed cylindrical passageways into plenum chamber 33. A particularadvantage of using tubes 32, as compared to merely perforations or holesof adequate diameter, lies in the formation of a gas pocket under plate31 in the areas around the individual tubes 32. We have found that thisis desirable because such a gas pocket trapped beneath tray or plate 31provides pressure surge dampening, which may result from flow changes ofthe mixture of hydrogen and liquid being supplied to the reactor vessel.However, the length of the tubes 32 is maintained as short as reasonablypossible to so function. Again, this is because of the desirability ofutilizing as little as possible of all processing space available invessel 11 for anything but contacting the feed streams with conversioncatalyst. A particular advantage to using tubes 32, as compared to acombination of tubes and perforations, is that the designed flowdistribution pattern is maintained over a wider range of flow rates.With tubes and perforations, gas normally flows up the perforations andliquid flows up the tubes. However, gas will find new flow paths throughthe tubes if the gas flow increases or the perforations become plugged,resulting in undesigned and potentially undesirable flow patterns.

To further assist in maintenance of plug-like flow of catalyst bed 10throughout its axial length, there is additionally provided in apreferred form or embodiment of the invention a plurality of axiallyspaced apart hydrogen gas redistribution or hydrogen gas-quenchingstages 39 within bed 10. In the arrangement of FIG. 1, the location ofone of the gas redistribution stages 39 is illustrated by the singleinverted angle member 40 extending transverse to the axis of bed 10. Thedetails of quench system 39 are best seen in FIGS. 5 to 7 where aplurality of inverted V-shaped sheds 40 (i.e. inverted angle members 40)are equally distributed over at least one transverse row extendinggenerally across the cross-sectional area of vessel 11. As shown in FIG.6 and in FIG. 7, a gas injection line 42 feeds an elongated tube 41extending through each individual shed 40 from a header 44 and branchlines 45 supplying the individual tubes 41. Desirably, but notnecessarily, a second tier of sheds 40 is axially spaced above the firsttier, with the sheds 40 in each tier being positioned at 90 degree(s) tothe other tier, as shown in FIG. 7. Construction of an individual shed40 is best seen in FIG. 6, wherein distribution pipe 41 includes aplurality of discharge holes 48, desirably proportioned to give equaldistribution of hydrogen gas along the full length of tube 41.Desirably, holes 48 are on the top side of tube 41 so that gas leavingthe tube is forced to flow downwardly within shed 40 to join gas risingfrom bed 10 under the area enclosed by the V-sides 49 of shed 40.Preferably, the full length of each skirt formed by sides 49 includesequally spaced slots 50 to exhaust both rising gas from bed 10 andquench gas entering from line 42. A particular value of the presentarrangement is that gas which may have become channeled in a portion ofthe bed below the quench system can be redistributed across the fullcross-sectional area of the bed to further avoid generation of local hotspots, eddy currents, or ebullation, within the upper portion of bed 10.

In accordance with another significant aspect of the present invention,FIG. 1 shows a catalyst replacement system, which in general comprises aseries of lock chambers for transferring fresh catalyst into bed 10through a pair of pressure lock chambers, including charging vessel 60and supply vessel 70. A similar series of lock chambers, includingdischarge vessel 80 and disposal vessel 90, transfer catalyst out of bed10. If necessary, a single pair of vessels could be used to charge anddischarge the catalyst, although the piping and sequencing procedurewould be more complex. In both cases, transfer flow is specificallydesigned to be as a liquid slurry and laminar to avoid undue abrasion ofcatalyst particles going into reactor vessel 11 and to avoid abruptagitation of the overlying bed of particles, with consequent ebullationand eddying of catalyst or fines in bed 10, when catalyst is withdrawnthrough inlet 30 of J-tube 29 at the bottom of reactor vessel 11.

To achieve laminar flow for supply of catalyst from charging vessel 60to the top of reactor vessel 11 or for catalyst removal from the bottomof bed 10 to discharge vessel 80, it is essential that the pressuredifferential between reactor vessel 11 and vessels 60 or 80, beaccurately controlled as by detecting the pressure differences betweensupply line 61 or discharge line 82 and reactor vessel 11. The pressuredifference is best zero when shut-off valves 64 or 84 are first openedor closed. The pressure differences between vessel 11 and line 61 ismeasured by gage 63 and pressure detectors 62 and 65. Differentialpressure gage 83 and detectors 81 and 85 serve a similar function tocontrol transfer of catalyst through valve 84 from the bottom of reactorvessel 11 to discharge vessel 80.

With reference particularly to supply of catalyst from vessel 60, itwill be understood, of course, that the vessel 60 is capable of beingbrought to a slightly higher pressure than the operating pressure ofreactor vessel 11, and closely controlled to assure that catalystsupplied to vessel 60 from supply vessel 70 is by laminar flow. For thispurpose, as indicated, vessels 70 and 60 are at atmospheric pressure,catalyst is first introduced into supply vessel 70 by way of funnel 100through line 101 and valve 102, and nitrogen is preferably flushedthrough supply vessel 70 through line 100 and/or line 71 to eliminateair and moisture that may be present on the catalyst. Either before orafter catalyst is introduced, vessel 70 is charged with a distillatehydrocarbon stream, preferably gas oil, to provide the necessaryslurrying liquid for mixing and transporting catalyst. This may eitherbe through funnel 100, valve 102, and line 101, or through line 104,valve 105 and line 106. Valve 102 is then closed and the catalyst isthen preferably heated to dehydrate and eliminate water from thecatalyst. It is to be understood that whenever the specification or theclaims states, mentions, or implies "mixing" or "admixing" or"commingling", or any of the like, including of any type(s) of catalyst,such stated, mentioned, or implied verbiage means within the spirit andscope of the present invention any type of "mixing" or "admixing" or"commingling", or any of the like, including any incidental mixing orany otherwise non-thorough/non-homogeneous mixing. Preferably, however,any type of "mixing" or "admixing" or "commingling", or any of the like,will be essentially thorough and/or essentially homogeneous.

An important requirement is that before transferring liquid to thecharging vessel 60, the pressure in supply vessel 70 must be equalizedto that in charging vessel 60, assuming, of course, that isolation valve64 between vessel 60 and the reactor vessel 11 is closed, and also thatvalves 67, 68 and 78 are closed. With valves 64, 67, 68, 78 and 102closed and pressure equalized between the vessels 60 and 70, transfervalve 75 may be opened to provide the same diameter path for thecatalyst slurry to flow throughout the path from J-tube 71 to vessel 60.The transfer is closely controlled by regulating the nitrogen gas flowrate and pressure introduced from line 104 through valve 105. Thepressure and flow rate are just sufficient to assure the desired laminarflow of catalyst into inlet 72 of J-tube 71 and thus upwardly throughline 76 and into charging vessel 60, which forms a catalyst chargingvessel. Laminar flow to transfer catalyst through J-tube 71 is entirelyin the liquid phase, with the catalyst as a slurry in the gas oil.Transfer of all catalyst is assisted by the funnel shape of bottom 79 ofvessel 70, and the position of intake 72 to J-tube 71 at the apex ofbottom 79. If all the catalyst in vessel 70 is transferred to vessel 60,flush oil from vessel 70 will naturally clear all the catalyst out ofline 76. However, to assure that all such catalyst has passed throughvalve 75 (so that valve 75 need not close on hard, abrasive catalystwith potential danger of scoring the valve 75 or the valve seat therein)additional flush fluid is preferably introduced from line 77 throughvalve 78 to clear line 76, either back into vessel 70, or forward intovessel 60.

With catalyst thus loaded into vessel 60, a similar procedure is usedfor transferring catalyst under laminar flow conditions as a liquidslurry into reactor vessel 11 through supply pipe 61 for distribution tothe top 20 of bed 10. If desired, of course, a deflector plate (notshown) may be used to distribute catalyst evenly across the top ofcatalyst bed 20. However, we have found that such a distribution aid isnot required. In the transfer of catalyst from the charging vessel 60 toreactor vessel 11, it will be understood that the pressure in vessel 60is brought to the pressure of reactor vessel 11. This is done byinjecting process hydrogen through valve 67. The oil should be heated toa temperature as close as possible to the temperature of reactants invessel 11, without vaporizing the oil. We have found this to beparticularly important to minimize any disturbance of thehydroprocessing process when fresh catalyst is added to an onstreamreactor vessel, such as reactor vessel 11. Once these requirements aremet, valve 64 should be opened for transfer. The actual laminar transferof the liquid slurry is controlled by valve 67 throttling the flow andpressure of hydrogen admitted from line 66. After transfer of thecatalyst, valve 68 in flush line 69 is opened briefly to assure that anycatalyst left in lines 61 and 19 is cleared before valve 64 is closed,for the reasons noted before. Excess hydrogen pressure in vessel 60 maybe relieved in a controlled manner via a suitable bleed line runningback to the common hydrogen source (not shown) of the hydroprocessingsystem.

Substantially continuous or intermittent transfer of deactivatedcatalyst for regeneration or disposal from the bottom of catalyst bed 10in reactor vessel 11 to discharge vessel 80 is controlled in the samemanner. As in all transfer of catalyst throughout the system of thepresent invention depicted in FIG. 1, the flow path from inlet 30 ofJ-tube 29, through line 82, including the bore of valve 84, is uniformin cross-sectional area and diameter. Similarly, transfer from dischargevessel 80 to disposal vessel 90 is through inlet 89 of J-tube 86 todischarge outlet 98 of line 92, including valve 94, into vessel 90.Deactivated catalyst is transferred laminarly from the bottom of thecatalyst bed 10 as a slurry in the hydrocarbon feed stream which, aspreviously mentioned, comprises the liquid hydrocarbon feed stream or amixture of hydrocarbon liquid feed and hydrogen-containing gas.Typically, the catalyst is transferred essentially in the liquidhydrocarbon feed stream (i.e. the liquid component of the hydrocarbonfeed stream).

In general the diameter of these laminar flow passageways are at leastfive times, and maybe as high as fifty or more times, the diameter ofthe individual particles to be passed therethrough. In this connectionto avoid jamming or obstruction, the inlets 72, 109, 30, 89 and 99 intotheir respective tubes 71, 108, 29, 86 and 96 are not flared orotherwise restricted, or perforated, so that all flow is solely anddirectly through the full and equal bore of such inlets. In the case ofcatalyst removal from reactor vessel 11, inlet 30 of tube 29 ispositioned at and over unperforated center plate 25 of catalyst supportscreen means 17, so that it is out of the direct flow of any hydrogengas stream rising through the innermost annular passageway formed bywalls 27 and radial ribs or spokes 26. This assures that flow into entry30 is substantially a liquid only slurry mixture with catalystparticles. Such a mixture at laminar flow conditions produces maximumcarrying capacity of the fluid. Additionally, the external dimensions ofthe circular bend or arc portion of the J-section of the tube 29 isseveral times the diameter of inlet 30 and the connected flow path,including the downwardly directed portion. The portion of tube 29 aboveinlet 30 is many times shorter and smaller in volume than the remainderof J-tube 29, down to, and including, control valve 84. A particularadvantage of keeping this portion of tube 29 small is to avoid thenecessity of forcing substantial amounts of catalyst back into the bed11 against the gravity head of catalyst bed 10 when that portion of theline is cleared at the end of each transfer.

Desirably, during periods when the catalyst is not being transferred, asmall amount of hydrogen may be continually bled through valve 88 intobed 10 through J-tube 29 to assure that catalyst particles do not clogentry 30. This avoids potential build up of coke at entry 30 of pipe 29.Such an arrangement assures that catalyst can be withdrawn by laminarflow without artificially fluidizing or levitating bed 10 directlyadjacent to J-tube entry 30.

Because gravity drainage of catalyst by an opening through the center ofthe catalyst support screen means 17 is not required in the presentarrangement, as in the prior art, it is possible to operate the entiresystem without use of solids handling valves. Accordingly, each of thetransfer valves in the present arrangement are preferably conventionalball valves formed with a single through bore in a rotatable ball.Specifically, we have found that conventional valves used to feed andcontrol flow of hydrocarbons, catalyst and hydrogen, into and out of thevessel 11, must seal against high pressure differentials between thevessel and the transfer vessels. For this service, a solid satellite,spherical-ball valve having a through bore of the same diameter as theinlet and outlet lines to the valve and metal-to-metal seals, providessuperior service when used in the catalyst transfer lines for carryingout the method of the present invention. Further, their commercial costand ready availability for such severity of service makes them mostuseful economically, both for initial installation and for servicereplacement. Valves manufactured by The Kaymr and Mogas Companies,called full-port valves are particularly useful in the presentembodiment. Further, the arrangement permits transfer of catalyst almostexclusively in a liquid phase which substantially reduces abrasion orcomminution of catalyst particles during transfer. Additionally,exclusion of entrained gas substantially improves the efficiency ofliquid transfer of catalyst particles and further reduces potentialdamage to the catalyst.

FIG. 4 illustrates a partial view of the bottom of pyramidal catalystbed support means 17 showing an alternate system for transferringcatalyst in a laminarly flowing liquid. In this embodiment, an L-valveis formed by vertical tube 54 and horizontal tube 52 for withdrawingcatalyst particles from the bottom of bed 10. As shown, intake 56 ispreferably directly above the central, non-perforated, section 25 of thetruncated pyramid formed by support means 17. While such an arrangementis less preferred than that shown in the embodiment of FIG. 1, such anarrangement is made suitable by the fact that the slurry of liquid andcatalyst can be made to flow only under laminar flow conditions. Witheither the J-tube of FIG. 1, or the L-valve of FIG. 4, arrangements, thepressure in discharge vessel 80 is brought up to equal that in reactorvessel 11. Valve 84 is opened and catalyst flow is controlled, as seenin FIG. 1, by regulating flow through valve 93. Such flow decreases thegas-pressure in discharge vessel 80 and line 82 sufficiently to induce alaminar flow of catalyst particles from vessel 11 when transfer valve 84is opened. After valve 84 has been flushed with vacuum gas oil throughvalve 88 and line 87 and then closed. The pressure in vessel 80 is thenreduced to a lower pressure (about 50 psig or less). The residuum isdrained from discharge vessel 80 through drain line 120, below J-tube 86and conical screen 121. Flush oil is then sent in through valve 93 towash residuum off the catalyst and to cool the catalyst. The dischargevessel 80 can be drained and filled as many times as needed. Thepressure in disposal vessel 90 is made equal to that in vessel 80 andvalve 94 is opened. The flow and pressure are then controlled throughvalve 110 to induce laminar flow of catalyst through J-tube 86 and intodisposal vessel 90. Valve 94 is flushed with flush oil through valve 107and closed. The contents of the disposal vessel 90 is preferably washedand cooled with flush oil which is then drained through drain line 122below conical screen 123. The spent catalyst contents of the disposalvessel 90 is then washed with water if desired through valve 110. Thedisposal vessel 90 should be purged of any hydrogen by sending innitrogen gas also through valve 110. Finally, disposal vessel 90 isnearly depressurized and the catalyst is dumped using water as thecarrier fluid through J-tube 96 by nitrogen flow through valve 110 tocontrol the rate of catalyst flow in discharge pipe 124.

continuing to refer to the drawings for other preferred embodiments ofthe present invention, a method is provided for maximally occupying areactor volume with a substantially packed bed of hydroprocessingcatalyst (e.g. catalyst bed 10) during hydroprocessing by contacting thesubstantially packed bed of hydroprocessing catalyst with an upflowinghydrocarbon feed stream having a liquid component and ahydrogen-containing gas component. The substantially packed bed ofhydroprocessing catalyst, preferably comprising the catalyst of thepresent invention as more particularly described below under thesubtitle "The Catalyst", is disposed in a reactor zone (or reaction zoneor zone for reaction) contained within a reactor volume (e.g. the entireinternal volumetric space available within the reactor vessel 11) suchthat the substantially packed bed of hydroprocessing catalyst occupiesat least about 50% by volume, preferably from about 75% by volume toabout 98% by volume of the reactor volume; more preferably from about80% by volume to about 98% by volume; most preferably at least about 90%by volume or from about 90% by volume to about 95% by volume of thereactor volume. Stated alternatively, hydroprocessing catalyst isdisposed or otherwise positioned within a reactor volume such thathydroprocessing catalyst occupies at least about 50% by volume,preferably from about 75% by volume to about 98% by volume of thereactor volume; more preferably from about 80% by volume to about 98% byvolume; most preferably at least about 90% by volume or from about 90%by volume to about 95% by volume of the reactor volume. "Reactor volume"(or the entire internal volumetric space available within the reactorvessel 11) means or may be generally defined as the volumetric spacewithin the reactor vessel 11 (or any similar hydroprocessing reactorvessel), including the summation or addition of the following internalvolumes: (i) an internal volume within the reactor vessel 11 representedby a volume (or internal cylindrical volume or main body volume of thereactor vessel 11) spanning or extending from an upper tangent line 180(see FIG. 8) to a lower tangent line 182 and generally illustrated asarrow TL in FIG. 8; and (ii) an internal volume within the upper domeclosure end 14 (or hemispherical head) of the reactor vessel 11; and(iii) an internal volume within the lower dome closure end 13 (orhemispherical bottom) of the reactor vessel 11. A "tangent line" isknown to those skilled in the art as a plane (i.e. horizontal plane)taken generally along the junctures of the sidewall 12 (which isessentially a straight upright wall) of the reactor vessel 11 with theupper and lower dome closure ends 14 and 13 respectively.

A hydroprocessing feed stream including a liquid component and ahydrogen-containing gas component upflows into the substantially packedbed of hydroprocessing catalyst at a rate of flow such that expansion ofthe substantially packed bed of hydroprocessing catalyst is limited toless than 10% by length beyond a substantially full axial length of thesubstantially packed bed of hydroprocessing catalyst in a packed bedstate. A volume of the hydroprocessing catalyst is withdrawn from thereactor zone to commence essentially plug-flowing downwardly of thesubstantially packed bed of hydroprocessing catalyst within the reactorzone; and hydroprocessing replacement catalyst is added to theessentially plug-flowing downwardly, substantially packed bed ofhydroprocessing catalyst at a rate to substantially replace the volumeof the withdrawn hydroprocessing catalyst. The procedure may be repeatedas many times as desired, even continuously repeated during continualhydroprocessing.

Another method is provided for hydroprocessing a hydrocarbon feed streamthat is upflowing through a hydroconversion reaction zone having asubstantially packed bed of catalyst which comprises forming a pluralityof annular mixture zones under a hydroconversion reaction zone having asubstantially packed bed of hydroprocessing catalyst such that each ofthe annular mixture zones contains a hydrocarbon feed stream having aliquid component and a hydrogen-containing gas component and wherein theannular mixture zones are concentric with respect to each other and arecoaxial with respect to the hydroconversion reaction zone. Thehydrocarbon feed stream from each of the annular mixture zones isintroduced into the substantially packed bed of hydroprocessing catalystto commence upflowing of the hydrocarbon feed stream from each of theannular mixture zones through the substantially packed bed of thecatalyst.

Considering the range of hydroconversion systems and/or hydroconversionreaction zones which could benefit from the preferred embodiments of thepresent invention, one skilled in the art will appreciate the variety ofcatalysts, having a variety of physical properties and elementalcompositions, which could be used in such a range of systems. It iswithin the spirit and scope of the present invention to encompass thesesystems employing catalysts having a size, shape and density which varywidely from system to system. However, it is important for the presentpreferred embodiment that the catalyst particles be of uniform and/orsame size, and shape (same density when in fresh catalyst state) withina single hydroconversion reaction zone of a hydroconversion system, inorder to achieve the desired catalyst and hydrocarbon flow patternswithin the hydroconversion reaction zone. A detailed description of thepreferred catalyst is presented below under the subtitle "THE CATALYST".It is to be understood that whenever the specification or the claimsstates, mentions, or implies "fresh catalyst", such stated, mentioned,or implied "fresh catalyst" means within the spirit and scope of thepresent invention any type of catalyst having any usable catalyst lifeor activity (e.g. regenerated catalyst, rejuvenated catalyst, partiallyfouled catalyst obtained from any source, etc.). Preferably, "freshcatalyst" means a type of catalyst that has never been used before andis obtained directly from a manufacturer with the lowest desired densityand the highest desired catalyst life or activity.

A hydroconversion system and/or a hydroconversion reaction zone of apresent preferred embodiment of the present invention contains acatalyst which is described in detail below under the subtitle "TheCatalyst", and may also be operated as a fixed bed (i.e. a catalyst bedwhich does not expand), a moving bed, an ebullated bed, an expanded bedor a fluidized bed configuration. A moving bed system is preferred.

By "moving bed", as used herein, is meant a reaction zone configurationin which a catalyst is added at one end of a catalyst bed in anintermittent or substantially continuous manner and is withdrawn at theother end in an intermittent or substantially continuous manner. A"moving bed" also includes a "plug-flow" or "plug flowing" catalyst bed10 or substantially packed bed of catalyst. As previously indicated,when any type of catalyst or catalyst bed 10 movement is mentioned,stated, or implied, the spirit and scope of the present inventionincludes such type of movement on any type of basis or in any manner(e.g. "periodic", "fully continuous", "non-continuous" etc.) , withoutthe necessity of having to specifically mention the type of basis ormanner. Preferably, catalyst is added at the top of the reaction zoneand withdrawn at the bottom. In the type of moving bed to which thepresent preferred embodiment is directed, the catalyst particles in thebed are substantially in contact with one another and plug-flowdownwardly. The catalyst bed is not significantly expanded when processfluids (e.g., liquid and gas) passes through it. It has essentially thecharacter of a fixed bed except for maybe a slight expansion upwardlyand for the addition and removal of catalyst. As the term is usedherein, a "moving bed" is not the same as a "fluidized bed", "ebullatingbed" or "expanded bed". In fluidized beds, the flow rate of a singlephase fluid, relative to the particles of the catalyst, is fast enoughso that the catalyst behaves like a fluid, with particles circulatingthroughout the bed or even being carried out of the bed with theproducts. Ebullating or expanded beds are very similar to fluidizedbeds, except that the relative rate of flow of two phase fluids (e.g.,liquid and gas) is regulated to expand the catalyst bed in random motionbetween 110% and 140% of the height of the catalyst in a "slumped" orpacked state. The typical ebullating bed reactor will have a mass ofsolid particles whose gross volume in the reaction vessel is at least 10percent larger when feed is flowing through it, as compared to thestationary mass with no feed flowing through it. Although the particlesin the bed do not necessarily circulate as if they were fluids, they areseparated from one another and go through random motion.

Several advantages ensue from use of a moving bed reactor. Byestablishing and maintaining appropriate gas and liquid velocities inpacked bed type reactors, just below the threshold of inertia that wouldcause the catalyst bed to fluidize and/or "channel" and/or lift thecatalyst into random motion, the uniform catalyst characteristicsdescribed above will allow the catalyst to migrate downward through thereactors in a predictable plugflow manner, as catalyst batches arewithdrawn from the reactor bottom. And further, by maintaining plug flowcatalyst movement downward within the reactors (e.g. reactor 11), thecatalyst within the reactors can be maintained in layers havingdiffering activity levels and reaction rates. The number of catalystlayers depend on the frequency of catalyst addition and withdrawal, andthe amount added and withdrawn in any given period of time. Typically,however, the number of different aged catalyst layers within the reactor(e.g. reactor 11) will be in the range from 10 to 60.

Intermittent or continuous catalyst additions and withdrawal may beused. Catalyst replacement rates can range from several percent of thecharge per day to several percent of the charge per week, depending onthe reactor size, catalyst metals loading capacity, feed rate, and feedcomposition and processing objectives. Fresh catalyst is introduced intothe downstream end of the catalyst bed (e.g. catalyst bed 10), and acorresponding volume of deactivated catalyst is removed from theupstream end of the catalyst bed, at a rate which is sufficient tomaintain the actual overall average level of catalytic upgradingactivity of the bed as a whole at or above the selected minimum averageactivity level. By "upstream" end of the catalyst bed (e.g. catalyst bed10), as used herein, is meant the end of the moving bed into which theheavy hydrocarbonaceous feed is introduced. By "downstream" end of thecatalyst bed is meant the end of the bed from which the process effluentis recovered. In a normal gravity flow system, the catalyst is added andeffluent removed at the top of the vessel (the downstream end). Spentcatalyst is withdrawn and feed introduced at the bottom (the upstreamend).

In a particularly important application of the present invention,catalyst is continuously added at the top of the reactor (e.g. reactor11) to the slowly moving bed (e.g. bed 10), and spent (and deactivatedcatalyst) catalyst is continuously withdrawn from bottom of the slowlymoving bed. The deactivated catalyst is removed from the reactor (e.g.reactor 11) after it has been deactivated to a substantially lower levelof activity than an acceptable minimum average level of activity of theoverall catalyst bed. This allows more efficient and complete use of thecatalyst activity, e.g. its metals capacity, for such feed upgradingfunctions as demetallation. As previously indicated, spent (anddeactivated) catalyst is withdrawn from the bottom of a reactor in ahydrocarbon liquid. One of the features of the present invention is thatthe hydrocarbon liquid that is withdrawing and transporting catalyst isthe liquid hydrocarbon component LH which is intended to flow upwardlythrough the bed of catalyst but has not. Thus, one of the features ofthe present invention is that the hydrocarbon liquid for transportingspent (and deactivated) catalyst is an unconverted liquid hydrocarboncomponent LH or a partially converted liquid hydrocarbon component LH ora mixture of both; and the transporting hydrocarbon liquid (i.e. theliquid hydrocarbon component LH) has not passed entirely upwardlythrough the catalyst bed.

The product from the method of the present invention exits a reactor(e.g. reactor vessel 11) and is normally subjected to furtherconventional refinery processing. All or part of the product can bepassed to a conventional, fixed bed upgrading operation, such as ahydrodesulfurization operation. Part of the product stream can berecycled, either for further catalytic treatment or as a diluent.Treatment of heavy feeds by catalytic demetallation according to thepresent process followed by fixed bed desulfurization is particularlyeffective, but all or part of a demetallized product from thecountercurrent demetallation reaction zone can also be processed in acountercurrent moving bed desulfurization reaction zone.

The present preferred embodiments of the present invention areapplicable to hydroconversion reaction zones for hydrocracking,hydrodemetallization, hydrotreating, hydrodesulfurization,hydrodenitrification, hydrofinishing and the like, all of whichcatalytically upgrade a heavy hydrocarbonaceous oil that represents theliquid hydrocarbon stream or liquid hydrocarbon feed stream (i.e. theliquid hydrocarbon component LH). By "heavy" liquid hydrocarbon stream,as used herein and as previously indicated, is meant liquid hydrocarbonstream at least 50 volume percent of which boils above about 204° C. andwhich preferably contains a substantial fraction boiling above about343° C. and particularly preferably above about 510° C. Preferred liquidhydrocarbon streams are residual fractions and synthetic crudes. Theycan be derived from crude petroleum, from coal, from oil shale, from tarsand bitumen, from heavy tar oils, and from other synthetic sources. Thepresent invention is advantageously employed to refine highly refractoryand contaminated liquid hydrocarbon streams. The liquid hydrocarbonstream may be substantially free from finely divided solids such asshale fines, sand or the like. Alternatively, the liquid hydrocarbonstream may contain a substantial concentration (e.g. about 1 weightpercent or more) of finely divided solids. As previously indicated, theliquid hydrocarbon stream (i.e. the liquid hydrocarbon component LH) ispreferably premixed with any type of hydrogen-containing gas (i.e. theliquid hydrocarbon component HG) which is preferably hydrogen, beforebeing introduced into the reactor vessel 11 as a single stream orhydrocarbon stream. The mixing ratios of the liquid hydrocarbon streamto the hydrocarbon containing gas may be any suitable ratio, well knownto those artisans possessing the ordinary skill in the art.

Typically, a heavy hydrocarbonaceous oil employed as a hydrocarbon feedstream in the present invention contains undesirable metals. Undesirablemetals which are often present in hydrocarbonaceous feeds notablyinclude nickel, vanadium, arsenic, and iron. These metals deactivateconventional, fixed bed catalysts (such as fixed bed hydroprocessingcatalysts) and also rapidly and irreversibly deactivate catalysts whenhigh metals level feed are charged directly to conventional units. Thesemetals are often present as organo-metallic compounds. Thus, the use ofthe terminology "iron, nickel, arsenic or vanadium compounds" means,those metals in any state in which they may be present in thehydrocarbon feed stream in the process of the present invention, eitheras metal particles, inorganic metal compounds, or an organo-metalliccompounds. Where amounts of metals are referred to herein, the amountsare given by weight based on the metal itself. For maximum efficiency insuch a countercurrent demetallation process, the hydrocarbon feed streamshould have levels of undesirable metals greater than about 150 ppm byweight of the hydrocarbon feed stream, preferably greater than about 200ppm by weight of the hydrocarbon feed stream, and more preferablygreater than about 400 ppm by weight. Although nickel, vanadium,arsenic, and iron are the usual metal contaminants, other undesiredmetals, such as sodium and calcium, may also contribute to the metalscontent of the hydrocarbon feed stream for purposes of catalyticdemetallation upgrading processing.

Catalytic upgrading conditions (e.g. catalytic desulfurizationconditions, catalytic hydrogenation conditions such as designed forasphaltenes saturation, catalytic denitrification conditions andcatalytic hydrocracking conditions, etc.) employed in thehydroconversions reaction zones within the reactor vessel 11 forpreferred embodiments of the present invention all include a reactiontemperature generally in the range of from about 230° C. to about 480°C., a pressure generally in the range of from about 30 to about 300atmospheres, a hydrogen rate ranging from about 1000 to about 10,000standard cubic feet per barrel of feed, and a liquid hourly spacevelocity (LHSV) in the range of from about 0.20 h-1 to about 10 h-1. Forfeed demetallation upgrading, the temperatures and pressures within thereaction zone can be those typical for conventional demetallationprocessing. The pressure is typically above about 500 psig (514.7 psia;35.5 bar). The temperature is typically greater than about 315° C., andpreferably above 371° C. Generally, the higher the temperature, thefaster the metals are removed; but the higher the temperature, the lessefficiently the metals loading capacity of the demetallation catalyst isused. While demetallation reaction can be conducted in the absence ofadded hydrogen, hydrogen is generally used and therefore requires fulland equal distribution into the moving bed along with any gases evolvingfrom the feed.

In carrying out a process of the preferred embodiments of the presentinvention, a minimum average level of catalytic feed upgrading activityfor the countercurrently moving catalyst bed (e.g. catalyst bed 10) as awhole is selected for the particular catalytic upgrading reaction. For amoving bed (e.g. catalyst bed 10) in a demetallation reaction system,for example, the minimum average upgrading activity level for thecatalyst bed is one which removes the necessary amount of metals fromthe hydrocarbon feed stream when it passes through the moving bed atdemetallation conditions. Similarly, for a desulfurization reactionsystem, the moving catalyst bed (e.g. catalyst bed 10) removes thenecessary amount of sulfur from the hydrocarbon feed stream when itpasses through the moving bed at desulfurization conditions. Thus, aswill be apparent to those skilled artisans, the minimum averageupgrading activity level for a particular reaction system will depend onthe desired degree of a contaminant, such as metals, sulfur, nitrogen,asphaltenes, etc., which the refiner desires to remove from the heavyoil feed. The degree of demetallation or desulfurization (or etc.) willtypically be set by economics and the downstream processing that theheavy feed will undergo. Further, according to preferred embodiments ofthe present invention, the actual average level of catalytic upgradingactivity for the catalyst bed (e.g. catalyst bed 10) as a whole ismeasured. Measurement of the actual average level of upgrading is madeby determining the extent to which the hydrocarbon feed stream is beingupgraded in the countercurrent moving bed system. For example, whenupgrading involves demetallation, demetallation activity is measured bya determination of the residual concentration of metals remaining in theliquid effluent stream from the moving bed system. When upgradinginvolves desulfurization, desulfurization activity is, analogously,measured by a determination of the residual concentration of sulfurremaining in the liquid effluent from the reaction system. Overallcatalyst bed upgrading activity for other upgrading reactions ismeasured in a similar manner by determining the residual amount of thecontainment which is to be removed by the process. In the presentpreferred embodiments, the rate at which catalyst is removed from thereaction zone, and the rate of catalyst replacement to the reactionzone, is established by a number of economic and operating factors,which include maintaining a desired average level of catalytic upgradingactivity.

THE CATALYST

In a preferred embodiment of the invention, the catalyst which ischarged to the reactor vessel 11 preferably satisfies the following fourmain criteria: (i) the catalyst has the appropriate catalytic activityand life for the particular application (e.g. demetallation,hydrodesulfurization, etc.); (ii) the catalyst has physical propertieswhich minimize its random motion in the reactor vessel 11; (iii) thecatalyst has physical properties which minimize catalyst loss both inthe catalyst transfer steps and in the reactor vessel 11; and (iv) thecatalyst is sufficiently uniform in size and shape and density toprevent classification by size in normal operation.

The catalyst in the present invention preferably has the appropriatecatalytic activity and life for the specific application (e.g.demetallation, hydrodesulfurization, etc.). For example, if the catalystis to be used for demetallation, it should have sufficient HDM activityand metals loading capacity (i.e. life) to meet the target demetallationwithout the use of uneconomic amounts of catalyst. The metals loadingcapacity of the catalyst is preferably greater than about 0.10 grams ofmetal per cubic centimeter of catalyst bulk volume and is morepreferably greater than about 0.20 grams metal per cubic centimeter ofcatalyst bulk volume. The catalyst properties which most affectcatalytic activity and metals loading capability are: pore structure(pore volume and pore size distribution); base material (e.g. aluminaversus silica); catalytic metals (amount, distribution, and type(nickel, molybdenum, cobalt, etc.)); surface area; and particle size andshape.

The catalyst in the present invention also preferably has physicalproperties which minimize catalyst lifting into random motion in theupflow type reactor vessel 11. Since one of the benefits of the presentinvention is the countercurrent contacting that is achieved between thereactants and catalyst, it is preferred to maintain plug flow of thecatalyst downwards through the entire length of the reactor vessel 11.The catalyst properties which are critical to minimizing or preventingcatalyst expansion are: catalyst particle density (highest particledensity possible is preferred while still meeting catalytic activity andmetals loading requirements); particle size (largest size practical ispreferred); skeletal density (higher skeletal density is preferred toreduce skeletal buoyancy); and size uniformity. One of the salientfeatures of the present invention is that the catalyst will not expandinto random motion in the reactor vessel 11, but will still move rathereasily during flow transportation. Under actual process conditionswithin the reactor, significantly smaller catalysts could rise to thetop while significantly larger catalysts could migrate to the bottom.This intervenes with optimal plug flow movement of catalyst. For thisreason, size specifications for the catalysts of the present inventionare narrower than those for fully packed or fixed bed and ebullated bedcatalysts.

The catalyst of the present invention should further have physicalproperties which minimize catalyst loss in the catalyst transfer stepsand in the reactor vessel 11. Breakage or attrition of the catalyst ineither the transfer steps or in the reactor vessel 11, can havesignificant adverse effects on the performance of the reactor systemitself and on any downstream equipment or processing unit. The followingcatalyst properties are critical to catalyst loss: catalyst attrition(minimum attrition is absolutely required); catalyst crush strength(maximum crush strength is required without producing a catalyst whichis very brittle and might suffer from excessive attrition); catalystsize and shape (spherical catalyst are preferred since they move moreeasily and have no rough or sharp edges to break off); and fines content(minimum fines is an absolute requirement to avoid adverse effects inthe reactor vessel 11 and downstream equipment).

The catalyst is sufficiently uniform in size and shape and density toprevent classification by size in normal operation. Generally, narrowspecifications are required for the catalyst to prevent classificationby size. Specific catalyst size is selected so that it is near the pointof being expanded into random motion, but not to the point of expansioninto random motion per se or ebullation.

All of the four main criteria for the selection of the catalyst of thepresent invention are important and are not independent or mutuallyexclusive of each other. The four main criteria must be balanced againsteach other to optimize the catalyst for the specific application. Forexample, to minimize catalyst expansion into random motion we wouldprefer a large and very dense catalyst. This is contrary to theproperties we might want for a residuum demetallation application wherewe need a small particle with low density diameters. These competingneeds must be balanced to ensure minimum catalyst expansion orebullation while achieving adequate catalytic activity and metalsloading capability, minimum attrition and minimum classification bysize.

Because there are competing catalyst requirements and because eachapplication is unique, the catalyst for the present invention may be anysuitable catalyst that is capable of assisting in the operation of theinvention and assisting in accomplishing the desired objects of theinvention.

The catalyst of the present invention unexpectedly produces aplug-flowing substantially packed bed (i.e. catalyst bed 10) ofhydroprocessing catalyst during hydroprocessing by contacting asubstantially packed bed of hydroprocessing catalyst with an hydrocarbonfeed stream (i.e. a liquid component and a hydrogen-containing gascomponent) that is upflowing at a rate controlled in an amount and to anextent sufficient to limit expansion of the substantially packed bed ofhydroprocessing catalyst to less than 10% by length beyond asubstantially full axial length of the substantially packed bed ofhydroprocessing catalyst in a packed bed state. More preferably, theexpansion of the substantially packed bed of hydroprocessing catalyst islimited to less than 5%, most preferably less than 2% or even less than1% , by length beyond a substantially full axial length of thesubstantially packed bed of hydroprocessing catalyst in a packed bedstate. The rate of flow of the hydrocarbon feed stream may be anysuitable rate controlled in an amount and to an extent sufficient tolimit the expansion of the substantially packed bed of hydroprocessingcatalyst, preferably the rate of flow is at a rate ranging from about0.01 ft/sec. to about 10.00 ft/sec.

The catalyst of the present invention more specifically unexpectedlyproduces a plug-flowing substantially packed bed of hydroprocessingcatalyst when a volume of the hydroprocessing catalyst is withdrawn ortransferred under preferably laminar flow conditions from the bottom ofthe substantially packed bed of hydroprocessing catalyst while, andsimultaneously to, the substantially packed bed of hydroprocessingcatalyst maximally and optimally occupies at least about 50% by volume,preferably at least about 75% by volume, preferably from about 80% byvolume to about 98% by volume (i.e. the entire internal and/or insideavailable volume) of the reactor vessel 11. The substantially packed bedof hydroprocessing catalyst of the present invention maximally andoptimally occupies a volume within the reactor vessel 11 that is largeror greater than a volume of a bed of catalyst in an ebullating reactorvessel that has substantially the same entire internal and/or insideavailable volume as the reactor vessel 11 and wherein the volume of thebed of catalyst in the ebullating reactor vessel is in a "slumped" (orpacked) catalyst bed condition or state. Typically, a bed of catalyst inan ebullating reactor vessel in a "slumped" catalyst bed conditionoccupies approximately up to less than about 75% by volume (maximum) ofthe entire internal and/or inside available volume of the ebullatingreactor vessel. Thus, the substantially packed bed of hydroprocessingcatalyst maximally and optimally occupies at least about 50% by volume,preferably at least about 75% by volume, preferably from about 80% byvolume to about 98% by volume of the entire internal and/or insideavailable volume of the reactor vessel 11. Most preferably, thesubstantially packed bed of hydroprocessing catalyst of the presentinvention maximally and optimally occupies from about 85% by volume toabout 95% by volume of the entire internal and/or inside availablevolume of the reactor vessel 11.

The catalyst of the present invention furthermore specificallyunexpectedly produces the plug-flowing substantially packed bed ofhydroprocessing catalyst when the volume of the hydroprocessing catalystis withdrawn or transferred in the hydrocarbon feed stream underpreferably laminar flow conditions from a central portion or section ofthe substantially packed bed of hydroprocessing catalyst and at alowermost or bottommost section thereof and below the entry points ofthe plurality of annular mixture zones MZ containing the hydrocarbonfeed stream (i.e. a liquid component and a hydrocarbon-containing gascomponent). As previously indicated, when the volume of thehydroprocessing catalyst of the present invention is withdrawn ortransferred to commence plug-flow, it is transferred or withdrawnpreferably laminarly in the liquid component of the hydrocarbon feedstream and is removed from above and in proximity to an impervious zone(i.e. imperforate center plate 25) of the bed support means 17 andsubstantially out of the flow path of the LH-HG mixtures (i.e. LH-HG₂,LH-HG₃, etc.) emanating out of the mixture zones MZ (i.e. MZ₂ s, MZ₃ s,etc.). The particular volume (or amount) of catalyst that is withdrawnat any desired time from the bottom of the substantially packed bed ofhydroprocessing catalyst may be any suitable volume or amount whichaccomplishes the desired objectives of the present invention.Preferably, such as by way of example only, the particular volume oramount of catalyst that is withdrawn at any desired time is a volume oramount ranging from about 0.10% by weight to about 25.00% by weight ofthe substantially packed bed (i.e. catalyst bed 10). The rate ofwithdrawal of a particular volume (or amount) of catalyst may also beany suitable volume or amount which accomplishes the desired objectivesof the present invention, such as a rate of withdrawal where the flowrate of the catalyst (e.g. the catalyst in the hydrocarbon feed stream)ranges from about 0.1 ft/sec. to about 20 ft/sec., more preferably fromabout 0.1 ft/sec. to about 10 ft/sec., and at a catalyst concentrationranging from about 0.10 lbs catalyst/lb. catalyst slurry (i.e. weight ofhydroprocessing catalyst plus weight of hydrocarbon feed stream) toabout 0.80 lbs catalyst/lb. catalyst slurry, more preferably from about0.15 lbs catalyst/lb. catalyst slurry to about 0.60 lbs catalyst/lb.catalyst slurry. As previously indicated, the withdrawn catalyst may beconveniently replaced by introducing a volume of fresh catalyst throughthe top of the reactor vessel 11 onto the catalyst bed 10. Thereplacement or catalyst addition rate may be any suitable replacement orcatalyst addition rate which will accomplish the desired objects of thepresent invention, such as a flow replacement rate of the replacementcatalyst (i.e. the replacement catalyst in the hydrocarbon refinedstream (e.g. gas oil)) ranging from about 0.1 ft/sec. to about 20ft/sec., more preferably from about 0.1 ft/sec. to about 10 ft/sec., andat a catalyst replacement concentration ranging from about 0.10 lbs.replacement catalyst/lb. catalyst slurry (i.e. weight of replacementcatalyst plus the hydrocarbon refined stream (e.g. gas oil) as theslurrying medium) to about 0.80 lbs replacement catalyst/lb. catalystslurry, more preferably from about 0.15 lbs catalyst/lb. catalyst slurryto about 0.60 lbs catalyst/lb. catalyst slurry.

In a preferred embodiment of the present invention, the catalyst of thepresent invention comprises an inorganic support which may includezeolites, inorganic oxides, such as silica, alumina, magnesia, titaniaand mixtures thereof, or any of the amorphous refractory inorganicoxides of Group II, III or IV elements, or compositions of the inorganicoxides. The inorganic support more preferably comprises a porous carriermaterial, such as alumina, silica, silica-alumina, or crystallinealuminosilicate. Deposited on and/or in the inorganic support or porouscarrier material is one or more metals or compounds of metals, such asoxides, where the metals are selected from the groups Ib, Vb, VIb, VIIb,and VIII of the Periodic System. Typical examples of these metals areiron, cobalt, nickel, tungsten, molybdenum, chromium, vanadium, copper,palladium, and platinum as well as combinations thereof. Preference isgiven to molybdenum, tungsten, nickel, and cobalt, and combinationsthereof. Suitable examples of catalyst of the preferred type comprisenickel-tungsten, nickel-molybdenum, cobalt-molybdenum ornickel-cobalt-molybdenum deposited on and/or in a porous inorganic oxideselected from the group consisting of silica, alumina, magnesia,titania, zirconia, thoria, boria or hafnia or compositions of theinorganic oxides, such as silica-alumina, silica-magnesia,alumina-magnesia and the like.

The catalyst of the present invention may further comprise additiveswhich alter the activity and/or metals loading characteristics of thecatalyst, such as but not limited to phosphorus and clays (includingpillared clays). Such additives may be present in any suitablequantities, depending on the particular application for thehydroconversion process including the applied catalyst. Typically, suchadditives would comprise essentially from about zero (0)% by weight toabout 10.0% by weight, calculated on the weight of the total catalyst(i.e. inorganic oxide support plus metal oxides).

Although the metal components (i.e. cobalt, molybdenum, etc.) may bepresent in any suitable amount, the catalyst of the present inventionpreferably comprises from about 0.1 to about 60 percent by weight ofmetal component(s) calculated on the weight of the total catalyst (i.e.inorganic oxide support plus metal oxides) or and more preferably offrom about 0.2 to about 40 percent by weight of the total catalyst, andmost preferably from about 0.5 to about 30 percent by weight of thetotal catalyst. The metals of Group VIII are generally applied in aminor or lesser quantity ranging from about 0.1 to about 30 percent byweight, more preferably from about 0.1 to about 10 percent by weight;and the metals of Group VIB are generally applied in a major or greaterquantity ranging from about 0.5 to about 50 percent by weight, morepreferably from about 0.5 to about 30 percent by weight; while aspreviously mentioned above, the total amount of metal components on theporous inorganic support is preferably up to about 60 percent by weight(more preferably up to about 40 percent by weight) of the totalcatalyst. The atomic ratio of the Group VIII and Group VIB metals mayvary within wide ranges, preferably from about 0.01 to about 15, morepreferably from about 0.05 to about 10, and most preferably from about0.1 to about 5. The atomic ratios would depend on the particularhydroprocessing application for the catalyst and/or on the processingobjectives.

The groups in the Periodic System referred to above are from thePeriodic Table of the Elements as published in Lange's Handbook ofChemistry (Twelfth Edition) edited by John A. Dean and copyrighted 1979by McGraw-Hill, Inc., or as published in The Condensed ChemicalDictionary (Tenth Edition) revised by Gessner G. Hawley and copyrighted1981 by Litton Educational Publishing Inc.

In a more preferred embodiment for the catalyst, the oxidichydrotreating catalyst or metal oxide component carried by or borne bythe inorganic support or porous carrier material is molybdenum oxide(MoO₃) or a combination of MoO₃ and nickel oxide (NiO) where the MoO₃ ispresent in the greater amount. The porous inorganic support is morepreferably alumina. The Mo is present on the catalyst inorganic support(alumina) in an amount ranging from about 0.5 to about 50 percent byweight, preferably from about 0.5 to about 30 percent by weight, mostpreferably from about 1.0 to about 20 percent by weight, based on thecombined weight of the inorganic support and metal oxide(s). When nickel(Ni) is present it will be in amounts ranging up to about 30 percent byweight, preferably from about 0.5 to about 20 percent by weight, morepreferably from about 0.5 to about 10 percent by weight, based on thecombined weight of the catalyst inorganic support and metal oxide(s).The oxidic hydrotreating catalyst or metal oxide component may beprepared by any suitable technique, such as by depositing aqueoussolutions of the metal oxide(s) on the porous inorganic supportmaterial, followed by drying and calcining. Catalyst preparativetechniques in general are conventional and well known and can includeimpregnation, mulling, co-precipitation and the like, followed bycalcination.

The catalyst has a surface area (such as measured by the B.E.T. method)sufficient to achieve the hydroprocessing objectives of the particularapplication. Surface area is typically from about 50 sq. meters per gramto about 300 sq. meters per gram, more typically from about 75 sq.meters per gram to about 150 sq. meters per gram.

The catalyst mean crush strength should be a minimum of about 5 lbs.Crush strength may be determined on a statistical sample of catalyticparticulates. For example, a fixed number (say 30 catalyst particles)are obtained from a statistical lot comprising a plurality of catalystparticles that are to be employed in the hydrogenation process of thepresent invention. Each catalyst particle is subsequently disposedbetween two horizontal and parallel steel plates. A force is thenapplied to the top steel plate until the disposed catalyst particlebreaks. The force applied to break the catalyst particle is the crushstrength. The test is repeated for the remaining catalyst particles, anda mean crush strength is obtained. Preferably further, no more thanabout 35% by wt. of the catalyst particles or catalytic particulateshave a mean crush strength of less than about 5 lbs.; more preferably,no more than about 15% by wt. of the catalyst particles or catalyticparticulates have a mean crush strength of less than about 5 lbs; andmost preferably, no more than about 0% by wt.

The catalyst of the present invention comprises a plurality of catalyticparticulates having a uniform size, which is preferably spherical with amean diameter having a value ranging from about 35 Tyler mesh to about 3Tyler mesh, more preferably ranging from about 20 Tyler mesh to about 4Tyler mesh, and most preferably from about 14 Tyler mesh to about 5Tyler mesh. The Tyler mesh designations referred to herein are from atable entitled "Tyler Standard Screen Scale Sieves" in the 1981 Editionof Handbook 53, published by CE Tyler Combustion Engineering, Inc., 50Washington St., South Norwalk, Conn. 06856.

Likewise, the preferred catalyst particle has a uniformly smooth androunded surface. Preferred shapes include, for example, spheres,spheroids, egg-shaped particles and the like. More preferably, thecatalyst of the present process is a rounded particle including aplurality of catalytic particulates having a size distribution such thatat least about 90% by weight of said catalytic particulates have anaspect ratio of less than about 2.0, more preferably equal to or lessthan about 1.5. As used herein, "aspect ratio" is a geometric termdefined by the value of the maximum projection of a catalyst particledivided by the value of the width of the catalyst particle. The "maximumprojection" is the maximum possible catalyst particle projection. Thisis sometimes called the maximum caliper dimension and is the largestdimension in the maximum cross-section of the catalyst particle. The"width" of a catalyst particle is the catalyst particle projectionperpendicular to the maximum projection and is the largest dimension ofthe catalyst particle perpendicular to the maximum projection.

The catalyst should have a particle size distribution such that thecatalyst bed 10 expands under the conditions within the reactor vessel11 to less than 10% by length (more preferably less than 5% and mostpreferably less than 1% by length) beyond a substantially full axiallength of the substantially packed bed of the hydroprocessing catalystin a packed bed state. In order to maximize reactor throughput, thecatalytic particulates have a narrow size distribution. The catalystemployed in the hydrogenation process of the present invention broadlycomprises a size range or size distribution such that at least about 90%by weight, preferably at least about 95% by weight, more preferably, atleast about 97% by weight, of the catalytic particulates in the catalystbed 10 have a diameter ranging from R₁ to R₂, wherein: (i) R₁ has avalue ranging from about 1/64 inch (i.e. the approximate opening size ofa 35 mesh Tyler screen) to about 1/4 inch (i.e. the approximate openingsize of a 3 mesh Tyler screen); (ii) R₂ also has a value ranging fromabout 1/64 inch (i.e. the approximate opening size of a 35 mesh Tylerscreen) to about 1/4 inch (i.e. the approximate opening size of a 3 meshTyler screen); and (iii) the ratio R₂ /R₁ has a value greater than orequal to about 1 and less than or equal to about 1.4 (or about thesquare root of 2.0). More preferably, the catalytic particulates in thecatalyst bed 10 have a diameter ranging from R₁ to R₂ wherein R₁ and R₂each has a value ranging from about 2/64 inch (i.e. the approximateopening size of a 20 mesh Tyler screen) to about 12/64 inch (i.e. theapproximate opening size of a 4 mesh Tyler screen), most preferably fromabout 3/64 inch (i.e. the approximate opening size of a 14 mesh Tylerscreen) to about 9/64 inch (i.e. the approximate opening size of a 5mesh Tyler screen), and wherein the ratio R₂ /R₁ has a value rangingfrom about 1.00 to about 1.4 (or about the square root of 2.0).

The catalyst employed in the hydrogenation process of the presentinvention also broadly comprises a size range or size distribution suchthat a maximum of about 2.0% by weight (more preferably a maximum ofabout 1.0% by weight and most preferably a maximum of about 0.5% byweight or less) of the catalyst particles or catalytic particulates hasa diameter less than R₁. The catalyst also has a size range or sizedistribution such that a maximum of about 0.4% by weight (morepreferably a maximum of about 0.2% by weight and most preferably amaximum of about 0.1% by weight or less) of the catalyst particles orcatalytic particulates have a diameter less than R₃, wherein R₃ is lessthan R₁ and the value of the ratio R₁ /R₃ is about 1.4 (or about thesquare root of 2.0). The catalyst particles or catalytic particulates ofthe catalyst preferably have a maximum attrition of about 1.0% by weight(more preferably a maximum of about 0.5% by weight and most preferably amaximum of about 0.25% by weight or less) of the catalyst particles orcatalytic particulates through a diameter (i.e., a Tyler screen opening)having a value of R₁, and a further maximum attrition of about 0.4% byweight (more preferably a maximum attrition of about 0.2% by weight andmost preferably a maximum attrition of about 0.1% by weight or less) ofthe catalyst particles or catalytic particulates through a diameter(i.e., again a Tyler screen opening) having a value of R₃ wherein R₃again (as stated above) is less than R₁ and the value of the ratio of R₁/R₃ is about 1.4 (or about the square root of 2.0). [Note that theattrition procedure is specified in ASTM D 4058-87. However, in thestandard method, the fines are removed through a 850μ (˜20 mesh) screen.In the present method, the screen is an opening equal to the minimumcatalyst size desired for the particular application, as morespecifically defined by the value of R₁ and R₃.] Thus, by way of exampleonly, for a catalyst with a specified size range of about 10 to about 12Tyler mesh, one would specify up to about 2.0% by wt. fines (morepreferably up to about 1.0% by wt.) MAX through 12 Tyler mesh and up toabout 0.4% by wt. (more preferably up to about 0.2% by wt.) MAX through14 Tyler mesh. Similarly, for a catalyst with a specified size range ofabout 6 to about 8 Tyler mesh, one would specify up to about 2.0% by wt.fines (more preferably up to about 1.0% by wt. fines) MAX through 8Tyler mesh and up to about 0.4% by wt. fines (more preferably up about0.2% by wt. fines) MAX through 10 Tyler mesh. For the catalyst with thespecified size range of about 10 to about 12 mesh, one would specify anattrition of up to about 1.0% by wt. (more preferably up to about 0.5%by wt., most preferably up to about 0.25% by wt.) MAX through 12 Tylermesh and up to about 0.4% by wt., (more preferably up to about 0.2% bywt., most preferably up to about 0.1% by wt.) MAX through 14 Tyler mesh.Similarly further, for catalyst with the specified size range of about 6to about 8 Tyler mesh, one would specify an attrition of up to about1.0% by wt. (more preferably up to about 0.5% by wt., most preferably upto about 0.25% by wt.) MAX through 8 Tyler mesh and up to about 0.4% bywt. (more preferably up to about 0.2% by wt., and most preferably up toabout 0.1% by wt.) MAX through 10 Tyler mesh.

The specific particle density of the catalyst particles is determined bythe requirements of the hydroconversion process. For the presentinvention it is preferred that the catalyst particles have a uniformdensity. By "uniform density" is meant that the density of at leastabout 70% by weight, preferably at least about 80% by weight, and morepreferably at least about 90% by weight, of the individual catalystparticles do not vary by more than about 10% from the mean density ofall catalyst particles; and more preferably the individual catalystparticles do not vary by more than about 5% from the mean density of allof the particles. In a preferred embodiment of the present invention thecatalyst (i.e. fresh catalyst) has a particle density ranging from about0.6 g/cc to about 1.5 g/cc, more preferably from about 0.7 g/cc to about1.2 g/cc, most preferably from about 0.8 g/cc to about 1.1 g/cc. Afterthe catalyst has at least been partially spent, the particle densitywould range from about 0.6 g/cc to about 3.0 g/cc, more preferably fromabout 0.7 g/cc to about 3.0 g/cc and most preferably from about 0.8 g/ccto about 3.0 g/cc. The particle size determination will remainsubstantially the same as defined above. Fines and attrition mayincrease during hydroprocessing.

While the catalyst of the present invention may be any catalyst asdefined above, we have discovered that the more preferred catalyst foroptimally accomplishing the objectives of the present inventioncomprises in combination the following properties: (i) a porousinorganic oxide support; (ii) one or more catalytic metals and/oradditional catalytic additives deposited in and/or on the porousinorganic oxide support; (iii) a crush strength at least about 5 poundsforce; (iv) a uniform size ranging from about 6 to about 8 Tyler meshsizes; (v) a fines content up to about 1.0 percent by weight through 8Tyler mesh and up to about 0.2 percent by weight through 10 Tyler mesh;(vi) an attrition up to about 0.5 percent by weight through 8 Tyler meshand up to of about 0.2 percent by weight through 10 Tyler mesh; (vii) agenerally uniform spherical shape; and (viii) a uniform density rangingfrom about 0.7 g/cc to about 3.0 g/cc. We have discovered unexpectedlythat the more preferred catalyst having or containing the immediateforegoing combination of properties, unexpectedly produces in an optimalfashion the plug-flowing substantially packed bed (i.e. catalytic bed11) of hydroprocessing catalyst which is simultaneously expanding toless than 10 percent by length (more preferably less than 1% by length)beyond a substantially full axial length of the substantially packed bedof hydroprocessing catalyst in a packed bed state while (andsimultaneously with) the substantially packed bed of hydroprocessingcatalyst maximally and optimally occupying from about 75 percent byvolume to about 98 percent by volume (i.e. the entire internal and/orinside available volume or reactor volume) of the reactor vessel

The particular type of porous base material or inorganic oxide support,the particular type of catalytic metal, the pore structure, the catalystsurface area and catalyst size, would all depend on the intendedspecific application (e.g. demetallation, desulfurization, etc.) of thecatalyst. Generally, the more preferred catalyst comprises a porousinorganic oxide support selected from the group consisting alumina,silica, and mixtures thereof, and has a surface area ranging from about75 square meters per gram to about 150 square meters per gram. Thepreferred catalyst comprises catalytic metal(s), present as oxide(s)deposited in and/or on the porous inorganic support. Oxide(s) of thecatalytic metal(s), or the metallic oxide component of the preferredcatalyst, is selected from the group consisting of molybdenum oxide,cobalt oxide, nickel oxide, tungsten oxide, and mixtures thereof, andcomprises from about 0.5 to about 50 percent by weight, more preferablyfrom about 0.5 to about 30 percent by weight, of the total catalyst(i.e. inorganic oxide support plus metal oxide(s)). The more preferredcatalyst further comprises a general uniform spherical shape having amean diameter ranging from about 20 Tyler mesh to about 4 Tyler mesh.While a spherical shaped catalyst is the more preferred catalyst, anextrudate may be employed if it is very strong, say having a crushstrength over 5 lbs. of force. The absolute size of the catalyst mayvary from application to application, but the more preferred catalysthas the narrow size distribution as previously stated above.

From the foregoing discussion it will be clear to the skilledpractitioner that, though the catalyst particles of the present processhave a uniform size, shape, and density, the chemical and metallurgicalnature of the catalyst may change, depending on processing objectivesand process conditions selected. For example, a catalyst selected for ademetallation application with minimum hydrocracking desired, could bequite different in nature from a catalyst selected if maximumhydrodesulfurization and hydrocracking are the processing objectives.The type of catalyst selected in accordance with and having theproperties mentioned above, is disposed in any hydroconversion reactionzone. A hydrocarbon feed stream is passed through the catalyst,preferably passed through such as upflow through the catalyst, in orderto hydroprocess the hydrocarbon feed stream. More preferably, thecatalyst is employed with the various embodiments of the presentinvention.

EXAMPLES

The following examples are exemplary examples of process runs, conductedin accordance with various method steps of the present invention andemploying the apparatus in accordance with various preferred embodimentsof the present invention. The following set-forth examples are given byway of illustration only and not by any limitation, and set-forth a bestmode of the invention as presently contemplated. All parameters such asconcentrations, flow rates, mixing proportions, temperatures, pressure,rates, compounds, etc., submitted in these examples are not to beconstrued to unduly limit the scope of the invention.

EXAMPLE I

In a semi-commercial scale residuum conversion pilot plant operating at100-200 BPD, the catalyst transfer procedure, as described above, wasdemonstrated more than 50 times. During each transfer, about 2 cubicfeet of catalyst was moved into and out of the reactor vessel runningcontinuously at typical residual desulfurization (RDS) conditions.Transfer rates up to 16 cubic feet per hour of catalyst wereaccomplished through pipes with an inside diameter 8 times larger thanthe catalyst diameter. Plug flow movement of the catalyst and theabsence of bed ebullation were proven using radioactively taggedcatalyst particles incorporated in the test bed.

Among the significant features of the invention specificallydemonstrated in such runs were that: (1) ball valves, such as those madeby the Kaymr and Mogas companies, can be used to isolate the RDS reactorfrom the catalyst transfer vessels, and to transfer catalyst particleswithout using solids handling valves, (2) the catalyst bed level andthus ebullation can be adequately monitored using a gamma-ray source anddetector, (3) J-tubes (all with upward flow sections substantiallyshorter than the downward flow paths) will satisfactorily transfercatalyst particles, without local ebullation, by laminar fluid flow, (4)use of fluid feed inlet distributor means with a conical support andconcentric annular segment plates prevents ebullation at the base of thecatalyst bed and provides adequate radial distribution of gas andliquid, by forming concentric alternate rings of gas and liquid (5)substantial differences were shown where bed ebullation (expansion)occurred with one catalyst as compared to no significant bed ebullation(expansion) with another catalyst using the same size and shape but witha lower density, and (6) transfer of catalyst into , and out of, a bedtravelling downwardly by gravity in a reactor vessel while continuouslyoperating a hydroprocessing system to react a gas containing hydrogenand a feed stream of hydrocarbon liquids flowing as a single stream fromthe bottom of the bed, will permit countercurrent flow withoutseparation during upward passage through and out of the top of thereactor vessel, and (7) as the result of intermittent catalystdischarge, the catalyst bed moves countercurrently down through thereactor in plug-like flow.

EXAMPLE II

From the foregoing tests in an apparatus under flow conditions describedin Example I, the effectiveness of the foregoing mechanical andhydraulic factors were validated for performing hydrotreating processingwith hydrocarbon and hydrogen streams counterflowing through a movingbed of catalyst particles, as follows:

In a pilot plant operating at up to 4 BPD hydrocarbon feed and hydrogenat 2200 PSI, catalyst bed expansion measurements were made atcommercial-scale flow velocities with beds of catalysts of differentsizes, shapes, and densities as indicated in Table I. Each type ofcatalyst was tested separately. Bed ebullation (expansion) was measuredusing a gamma-ray source and detector means mounted to detect 10% bedexpansion. Table I shows flow velocities required to produce 10% bedexpansion with several catalysts at a standard hydrogen recirculationrate of 5000 SCFB. These results confirmed the bed expansion resultsfrom the semi-commercial scale plant of Example I.

Table II is a similar set of runs using beds of three of the samecatalyst particles as those tested under conditions shown in Table Iexcept that the liquid viscosity, liquid density and pressure of thehydrocarbon feed stream and gas were lower in Table II than Table I tomatch a different set of commercial operating conditions. From Tables Iand II the effect of catalyst particle size, density and shape areclearly indicated for different flow conditions for the liquid and gascomponents of the feed. The design feed rates for a hydrocarbon treatingprocess were calculated by standard scaling procedures to indicate thevalues in MBPD (thousands of barrels per day) through a reactor vesselcontaining a catalyst bed 11.5 feet in diameter.

In general catalyst for commercial use would be selected on the basis oflevitation or ebullation at a selected rate which is substantiallyhigher than normal design feed rate, say up to 100% greater.Additionally, these tests indicate that some commercial catalysts willnot lift at reasonable design feed rates if the particles have a highdegree of uniformity and are sufficiently strong to maintain theirintegrity during movement into and out of the reactor vessel, withoutattrition or breakage.

                                      TABLE I                                     __________________________________________________________________________    CATALYST BED EXPANSION STUDY TEST RESULTS                                     2200 PSI Hydrogen and Flush Oil                                               Liquid Density 51 lb/ft 3 Viscosity 1.1 cp                                    Gas Density 0.49 lb/ft 3 Viscosity 0.016 cp                                                               Flow rates for 10% Bed Expansion                                              @ 5000 SCFB H.sub.2                               Cata-                                                                              Relative    Skeletal                                                                           Particle                                                                            Effective Density                                                                      Liquid Velocity                                                                        Gas Velocity                                                                          MBPD in 11.5            lyst Size Shape  Density                                                                            Density                                                                             In Oil.sup.(1)                                                                         Ft/Min   Ft/Sec  Ft. ID                  __________________________________________________________________________                                                          Reactor                 A    1    Cylinder                                                                             2.69 1.05  0.55     0.46     0.11    13                      B    1.6  Quadralobe                                                                           3.55 1.03  0.56     0.60     0.14    17                      C    2    Cylinder                                                                             3.61 1.60  1.05     0.46     0.11    13                      D    3.2  Sphere 2.33 0.60  0.21     0.32     0.07    9                       E    3.2  Sphere 3.63 0.83  0.47     1.38     0.33    40                      F    3.2  Cylinder                                                                             3.58 1.37  0.89     1.38     0.33    40                      __________________________________________________________________________     .sup.(1) Effective Density in Oil = Density of the Particle in Oil with       Buoyancy Forces Includes = (Skeletal Density) (vol % Skeleton) + (Oil         Density) (vol % Pores)  Oil Density                                      

                                      TABLE II                                    __________________________________________________________________________    CATALYST BED EXPANSION STUDY TEST RESULTS                                     With Hydrogen and Hydrocarbon at 1000 PSI                                     Liquid Density 48 lb/ft 3 Viscosity 0.56 cp                                   Gas Density 0.23 lb/ft 3 Viscosity 0.013. cp                                                              Flow rates for 10% Bed Expansion                                              @ 5000 SCFB H.sub.2                               Cata-                                                                              Relative    Skeletal                                                                           Particle                                                                            Effective Density                                                                      Liquid Velocity                                                                        Gas Velocity                                                                          MBPD in 11.5            lyst Size Shape  Density                                                                            Density                                                                             In Oil.sup.(1)                                                                         Ft/Min   Ft/Sec  Ft. ID                  __________________________________________________________________________                                                          Reactor                 C    2    Cylinder                                                                             3.61 1.60           0.53     0.13    15                      E    3.2  Sphere 3.63 0.83           1.38     0.33    40                      F    3.2  Cylinder                                                                             3.58 1.37           1.50     0.50    60                      __________________________________________________________________________

EXAMPLE III

In a 4 foot diameter vessel a "cold model" was operated using up to 8000BPD water and 275 SCFM air. The features of the inlet liquid and gasdistributor as well as the hydrogen gas redistribution and quenchstages, described above and shown in the drawings were scaled andtested. Flow measurements and underwater photography proved thatdistribution of the inlet gas and liquid was uniform across the fullcross-sectional area of the catalyst support screen in the vessel.Redistribution of the rising gas through the inverted V-shaped sheds wasshown to be surprisingly effective even when gas was intentionallymaldistributed below the redistributor stages.

SUMMARY OF TEST RESULTS FOR EXAMPLES I,II AND III

Briefly, these test results show that the present invention makespossible substantially continuous flow of uniformly distributed hydrogenand hydrocarbon liquid across a densely packed catalyst bed to fillsubstantially the entire volume of a reactor vessel by introducing thefluids as alternate annular rings of gas and liquid at a rateinsufficient to levitate the bed and with the catalyst selected with adensity, shape and size adequate to prevent lifting of the bed at thedesired feed rates. (Catalysts are selected by measuring bed expansionin a large pilot plant run with hydrocarbon, hydrogen, and catalyst atthe design pressures and flow velocities). At the desired flow rate,such catalyst continually flows in a plug-like manner downwardly throughthe vessel by introducing fresh catalyst at the top of the bed bylaminarly flowing such catalyst in a liquid stream on a periodic orsemicontinuous basis. Catalyst is removed by laminarly flowing catalystparticles in a liquid stream out of the bottom of the bed. Intake forsuch flow is out of direct contact with the stream of gas flowingthrough the bed and the flow path is substantially constant incross-sectional area and greater in diameter by several times than thediameter of the catalyst particles.

EXAMPLE IV

A plurality of catalytic particulates were charged into a reaction zonecontained within a reactor, such as reactor vessel 11. The plurality ofcatalytic particulates formed a catalyst bed (such as catalyst bed 10 inFIGS. 1, 8 and 9). The catalyst bed was supported in the reactor by atruncated conical bed support similar to the support that is generallyillustrated as 17 in FIGS. 8, 9 and 11-13. An inlet distributor, such ascircular plate member 31 in FIGS. 1 and 11 with the multiplicity oftubes 32, extended across a full cross-sectional area of the reactorunderneath the truncated conical bed support to form a plenum or inletchamber between the inlet distributor and the truncated conical bedsupport, as generally illustrated in FIGS. 8, 9 and 11-13. The truncatedconical bed support for the catalyst bed included a series of annularpolygons that included a plurality of segmented plates (such assegmented plates 27 in FIGS. 2 and 3) connected to or formed with radialspoke members such as members 26 FIGS. 10-13. The plurality of segmentedplates, each having a thickness of about 10 inches and a width of about1.5 inch, were secured to 8 radial spoke members. The interengagedsegmented plates and radial spoke members formed a web-like structure toproduce essentially annularly continuous mixture zones for receiving aflow of hydrocarbon feed stream, and were overlayed with a screen havingscreen openings with a mean diameter that was smaller than the catalyticparticulates. Each mixture zone underneath the screen had a generallycircumferentially uniform thickness.

The catalytic particulates comprised an alumina porous carrier materialor alumina inorganic support. Deposited on and/or in the alumina porouscarrier material was an oxidic hydrotreating catalyst componentconsisting of NiO and/or MoO₃. The Mo was present on and/or in thealumina porous carrier material in an amount of about 3% by wt., basedon the combined weight of the alumina porous carrier material and theoxidic hydrotreating catalyst component(s). The Ni was present on and/orin the alumina porous carrier material in an amount of about 1% by wt.,based on the combined weight of the alumina porous carrier material andthe oxidic hydrotreating catalyst component(s). The surface area of thecatalytic particulates was about 120 sq. meters per gram.

The plurality of catalytic particulates were generally spherical with amean diameter having a value ranging from about 6 Tyler mesh to about 8Tyler mesh and an aspect ratio of about 1. The mean crush strength ofthe catalytic particulates was about 5 lbs. force. The metals loadingcapacity of the catalyst or plurality of catalytic particulates wasabout 0.3 grams of metal per cubic centimeter of catalytic particulatebulk volume.

The catalytic particulates had a size distribution such that 98.5% byweight of the catalytic particulates in the catalyst bed had a diameterranging from R₁ to R₂ wherein: (i) R₁ had a value of about 0.093 inch(i.e. the approximate opening of an 8 mesh Tyler screen); (ii) R₂ had avalue of about 0.131 inch (i.e. the approximate opening size of a 6 meshTyler screen); and (iii) the ratio R₂ /R₁ had a value equal to about thesquare root of 2.0 or about 1,414. The size distribution of thecatalytic particulates was also such that a maximum of about 1.0% byweight of the catalytic particulates had a diameter less than R₁. Thecatalyst further also had a size distribution such that a maximum ofabout 0.2% by weight of the catalytic particulates had a diameter lessthan R₃, wherein R₃ was less than R₁ and the value of the ratio R₁ /R₃was about the square root of 2.0 or about 1.414.

The catalytic particulates of the catalyst had a maximum attrition ofabout 0.5% by weight of the catalytic particulates through a diameter(i.e. a Tyler screen opening) having the value of R₁, and a furthermaximum attrition of about 0.2% by weight of the catalytic particulatesthrough a diameter (i.e. a Tyler screen opening) having the value of R₃wherein R₃ again was less than R₁ and the value of the ratio of R₁ /R₃was about the square root of 2.0 or about 1.414. Stated alternatively,for the catalytic particulates with the specified size range ordistribution of about 6 to about 8 Tyler mesh, the specified attritionfor the catalytic particulates was up to about 0.5% by weight MAXthrough 8 Tyler mesh and up to about 0.2% by weight MAX through 10 Tylermesh.

The catalytic particulates had a maximum fines content of up to about1.0% by wt. through 8 Tyler mesh and up to about 0.2% by wt. through 10Tyler mesh. Stated alternatively, for the catalytic particulates withthe specified size range or distribution of about 6 to about 8 Tylermesh, the specified fines content for the catalytic particulates was upto about 1.0% wt. fines MAX through 8 Tyler mesh and up to about 0.2% bywt. fines MAX through 10 Tyler mesh. The catalytic particulates had auniform density such that mean density of the catalytic particulateswere about 0.9 g/cc.

The liquid component of the hydrocarbon feed stream was a heavyatmospheric residuum wherein at least 95% by volume of which boiledabove about 343° C. and wherein a substantial fraction (e.g. 50% byvolume) boiled above about 510° C. The "heavy" hydrocarbon feed had anundesirable metal content of about 90 ppm by weight of the "heavy"hydrocarbon feed. The hydrogen-containing gas of the hydrocarbon feedstream was essentially 97% pure hydrogen and was mixed with the heavyatmospheric residuum stream in a mixing ratio of 623 liters ofhydrogen-containing gas at standard conditions per liter of heavyatmospheric residuum in order to form the hydrocarbon feed stream.

The hydrocarbon feed stream was passed through the inlet distributor andintroduced into the plenum chamber of reactor at a flow rate rangingfrom about 0.1 ft/sec. to about 1.00 ft/sec. The hydroprocessingpressure and temperature within the reactor were about 2300 psig. andabout 400° C. respectively. From the plenum chamber of the reactor thehydrocarbon feed stream entered into the annular continuous mixturezones and was uniformly fed through the screen and into the catalyst bedsuch as not to induce local ebullation or eddy currents in the catalystbed, especially in proximity to the conical bed support which wasoverlayed with the screen.

The catalyst bed in the reactor contained a plurality of axially spacedapart hydrogen gas redistribution (or hydrogen gas-quenching) assemblies(see FIGS. 5 and 7 as illustrative of the hydrogen gas-quenchingassemblies). As the hydrocarbon feed stream flowed upwardly through thecatalyst bed, hydrogen gas was emitted from the hydrogen gasredistribution assemblies, which redistributed any hydrogen-containinggas that had become channeled in a portion of the catalyst bed below (orin close proximity to) the hydrogen gas redistribution assemblies andfurther avoided generation of local hot spots, eddy currents orebullation in the upper part (especially above the hydrogen gasredistribution assemblies) of the catalyst bed.

The liquid hydrocarbon feed stream exited the reactor at a withdrawalflow rate of about 3.6 ft/sec. and had been upgraded such that itcontained a metal content of about 3 ppm by wt. of the liquidhydrocarbon feed stream. As the hydrocarbon feed stream flowed upwardlythrough the catalyst bed, a gamma ray source in the catalyst bed incombination with a gamma ray detector on the reactor (see for examplegamma ray source 22 in the catalyst bed 10 competing with the gamma raydetector 24 on the reactor vessel 10 in FIG. 1) detected that thecatalyst bed expanded less than 10% by length over or beyondsubstantially the full axial length of the catalyst bed in a packed bedstate.

After the reactor was on stream for about 1 weeks, approximately 7.25cubic meters (or about 3.3% by weight of the catalyst bed) of catalyticparticulates were laminarly withdrawn in the hydrocarbon feed streamthrough a J-tube (such as J-tube 29 in FIG. 1) at a flow rate of about3.6 ft/sec. The withdrawn catalyst in the hydrocarbon feed stream had aconcentration of about 0.5 lbs. catalyst/lb. catalyst slurry (i.e.weight of withdrawn catalyst plus weight of hydrocarbon feed stream).When and/or as the volume of catalytic particulates were withdrawn ortransferred from the bottom of the catalyst bed, the catalyst bed (i.e.a substantially packed bed of catalyst) began to plug-flow. Thewithdrawn catalyst was replaced by introducing a comparable volume offresh replacement catalyst through the top of the reactor. The freshreplacement catalyst was slurried in a hydrocarbon refined stream (e.g.gas oil) and was introduced into the reactor at a flow catalystreplacement rate of about 3.6 ft/sec., and at a catalyst replacementconcentration of about 0.5 lbs. replacement catalyst/lb. catalyst slurry(i.e. weight of replacement catalyst plus the hydrocarbon refined stream(e.g. gas oil) as the slurrying medium).

While the present invention has been described herein with reference toparticular embodiments thereof, a latitude of modification, variouschanges and substitutions are intended in the foregoing disclosure, andit will be appreciated that in some instances some features of theinvention will be employed without a corresponding use of other featureswithout departing from the scope of the invention as set forth.

We claim:
 1. A method for producing an essentially downwardlyplug-flowing substantially packed bed of hydroprocessing catalyst withina hydroconversion reaction zone comprising the steps of:(a) forming aplurality of annular mixture zones under a hydroconversion reaction zonehaving a substantially packed bed of hydroprocessing catalyst such thateach of said annular mixture zones contains a hydrocarbon feed streamhaving a liquid component and a hydrogen-containing gas component andwherein said annular mixture zones are concentric with respect to eachother and are coaxial with respect to said hydroconversion reactionzone; (b) introducing said hydrocarbon feed stream from each of saidannular mixture zones of step (a) into said substantially packed bed ofhydroprocessing catalyst to commence upflowing of said hydrocarbon feedstream from each of said annular mixture zones through saidsubstantially packed bed of the catalyst; and (c) withdrawing a volumeof particulate catalyst from said hydroconversion reaction zone toproduce an essentially downwardly plug-flowing substantially packed bedof hydroprocessing catalyst within said hydroconversion reaction zone.2. The method of claim 1 wherein said hydroprocessing catalyst comprisesa plurality of catalytic particulates having a mean diameter rangingfrom about 35 Tyler mesh to about 3 Tyler mesh; and a size distributionsuch that at least about 90% by weight of said catalytic particulateshave a diameter ranging from R₁ to R₂, wherein:(1) R₁ has a valueranging from about 1/64 inch to about 1/4 inch, (2) R₂ has a valueranging from about 1/64 inch to about 1/4 inch, (3) a value of a ratioR₂ /R₁ ranges from about 1.0 to about 1.4; andan aspect ratio of lessthan about 2.0.
 3. The method of claim 2 wherein said catalyticparticulates have a size distribution such that at least about 95% byweight of said catalytic particulates have a diameter ranging from R₁ toR₂.
 4. The method of claim 2 wherein said catalytic particulates have asize distribution such that at least about 97% by weight of saidcatalytic particulates have a diameter ranging from R₁ to R₂.
 5. Themethod of claim 2 wherein said catalytic particulates have a sizedistribution such that a maximum of about 2.0% by weight of saidcatalytic particulates have a diameter less than R₁.
 6. The method ofclaim 2 wherein said catalytic particulates have a size distributionsuch that a maximum of about 1.0% by weight of said catalyticparticulates have a diameter less than R₁.
 7. The method of claim 1additionally comprising disposing a plurality of inert pellets undersaid reaction zone prior to said introducing step (b).
 8. The method ofclaim 2 wherein each of said plurality of annular mixture zones isessentially an annularly continuous mixture zone.
 9. The method of claim8 wherein each of said plurality of annular mixture zones has agenerally uniform thickness.
 10. The method of claim 2 wherein saidcatalytic particulates comprise alumina carrying a metal oxide selectedfrom the group consisting of molybdenum oxide, nickel oxide, andmixtures thereof.
 11. The method of claim 8 wherein said catalyticparticulates comprise alumina carrying a metal oxide selected from thegroup consisting of molybdenum oxide, nickel oxide, and mixturesthereof.
 12. The method of claim 10 wherein said catalytic particulateshave a minimum mean crush strength of about 5 lbs.
 13. The method ofclaim 12 wherein said aspect ratio is equal to or less than about 1.5.14. The method of claim 13 wherein said introducing step (b) comprisesupflowing through said catalyst of step (a) a hydrocarbon feed stream ata hate of flow such that said catalytic particulates expand to less than5% by length beyond a substantially full axial length of the catalyticparticulates in a packed bed state.
 15. The method of claim 13 whereinsaid introducing step (b) comprises upflowing through said catalyst ofstep (a) a hydrocarbon feed stream at a rate of flow such that saidcatalytic particulates expand to less than 1% by length beyond asubstantially full axial length of the catalytic particulates in apacked bed state.
 16. The method of claim 7 additionally comprisingdisposing a plate member under said hydroconversion reaction zone suchthat said plate member is totally spaced therefrom, said plate membercomprising a multiplicity of tubes bound thereto and axially extendingdownwardly therefrom for receiving said hydrocarbon feed stream and forconducting the same into a plenum chamber positioned under saidhydroconversion reaction zone; and disposing a permeable screen on saidplate member for supporting said inert pellets.
 17. A method forhydroprocessing a hydrocarbon feed stream that is upflowing through ahydroconversion reaction zone having a substantially packed bed ofcatalyst comprising the steps of:(a) forming a plurality of annularmixture zones under a hydroconversion reaction zone having asubstantially packed bed of hydroprocessing catalyst such that each ofsaid annular mixture zones contains a hydrocarbon feed stream having aliquid component and a hydrogen-containing gas component and whereinsaid annular mixture zones are concentric with respect to each other andare coaxial with respect to said hydroconversion reaction zone, andwherein said hydroprocessing catalyst comprises a plurality of catalyticparticulates having a mean diameter ranging from about 35 Tyler mesh toabout 3 Tyler mesh; and a size distribution such that at least about 90%by weight of said catalytic particulates have an aspect ratio of lessthan about 2.0 and a diameter ranging from R₁ to R₂, wherein:(1) R₁ hasa value ranging from about 1/64 inch to about 1/4 inch, (2) R₂ has avalue ranging from about 1/64 inch to about 1/4 inch; (3) a value of aratio R₂ /R₁ ranges from about 1.0 to about 1.4; and (b) introducingsaid hydrocarbon feed stream from each of said annular mixture zones ofstep (a) into said substantially packed bed of hydroprocessing catalystto commence upflowing of said hydrocarbon feed stream from each of saidannular mixture zones through said substantially packed bed of thecatalyst.
 18. The method of claim 17 wherein said catalytic particulateshave a size distribution such that at least about 95% by weight of saidcatalytic particulates have a diameter ranging from R₁ to R₂.
 19. Themethod of claim 17 wherein said catalytic particulates have a sizedistribution such that at least about 97% by weight of said catalyticparticulates have a diameter ranging from R₁ to R₂.
 20. The method ofclaim 17 wherein said catalytic particulates have a size distributionsuch that a maximum of about 2.0% by weight of said catalyticparticulates have a diameter less than R₁.
 21. The method of claim 17wherein said catalytic particulates have a size distribution such that amaximum of about 1.0% by weight of said catalytic particulates have adiameter less than R₁.
 22. The method of claim 17 additionallycomprising disposing a plurality of inert pellets under said reactionzone prior to said introducing step (b).
 23. The method of claim 17wherein each of said plurality of annular mixture zones is essentiallyan annularly continuous mixture zone.
 24. The method of claim 23 whereineach of said plurality of annular mixture zones has a generally uniformthickness.
 25. The method of claim 17 wherein said catalyticparticulates comprise alumina carrying a metal oxide selected from thegroup consisting of molybdenum oxide, nickel oxide, and mixturesthereof.
 26. The method of claim 22 wherein said catalytic particulatescomprise alumina carrying a metal oxide selected from the groupconsisting of molybdenum oxide, nickel oxide, and mixtures thereof. 27.The method of claim 25 wherein said catalytic particulates have aminimum mean crush strength of about 5 lbs.
 28. The method of claim 27wherein said aspect ratio is equal to or less than about 1.5.
 29. Themethod of claim 28 wherein said introducing step (b) comprises upflowingthrough said catalyst of step (a) a hydrocarbon feed stream at a rate offlow such that said catalytic particulates expand to less than 5% bylength beyond a substantially full axial length of the catalyticparticulates in a packed bed state.
 30. The method of claim 28 whereinsaid introducing step (b) comprises upflowing through said catalyst ofstep (a) a hydrocarbon feed stream at a rate of flow such that saidcatalytic particulates expand to less than 1% by length beyond asubstantially full axial length of the catalytic particulates in apacked bed state.
 31. The method of claim 22 additionally comprisingdisposing a plate member under said hydroconversion reaction zone suchthat said plate member is totally spaced therefrom, said plate membercomprising a multiplicity of tubes bound thereto and axially extendingdownwardly therefrom for receiving said hydrocarbon feed stream and forconducting the same into a plenum chamber positioned under saidhydroconversion reaction zone; and disposing a permeable screen on saidplate member for supporting said inert pellets.
 32. A method Tonhydroprocessing a hydrocarbon feed stream that is upflowing through ahydroconversion reaction zone having a substantially packed bed ofcatalyst comprising the steps of:(a) forming a plurality of annularmixture zones under a hydroconversion reaction zone having asubstantially packed bed of hydroprocessing catalyst such that each ofsaid annular mixture zones contains a hydrocarbon feed stream having aliquid component and a hydrogen-containing gas component and whereinsaid annular mixture zones are concentric with respect to each other andare coaxial with respect to said hydroconversion reaction zone; and (b)introducing said hydrocarbon feed stream from each of said annularmixture zones of step (a) into said substantially packed bed ofhydroprocessing catalyst to commence upflowing of said hydrocarbon feedstream from each of said annular mixture zones through saidsubstantially packed bed of the catalyst.
 33. The method of claim 32wherein said step (a) forming a plurality of annular mixture zones undera hydroconversion reaction zone having a substantially packed bed ofhydroprocessing catalyst additionally comprises forming said pluralityof annular mixture zones with at least one means for reducing a size ofa hydrogen-containing gas bubble from said hydrogen-containing gascomponent of said hydrocarbon feed stream.
 34. The method of claim 32wherein said step (a) forming a plurality of annular mixture zones undera hydroconversion reaction zone having a substantially packed bed ofhydroprocessing catalyst additionally comprises forming each of saidplurality of annular mixture zones to comprise a generally uniformthickness ranging from about 1 inch to about 4 feet.
 35. The method ofclaim 32 wherein said step (b) introducing said hydrocarbon feed streamfrom each of said annular mixture zones of step (a) into saidsubstantially packed bed of hydroprocessing catalyst comprises flowingupwardly said hydrocarbon feed stream from each of said annular mixturezones of step (a) into said substantially packed bed of hydroprocessingcatalyst at a rate of flow such that said substantially packed bed ofhydroprocessing catalyst expands to less than 10% by length beyond asubstantially full axial length of said substantially packed bed ofhydroprocessing catalyst in a packed bed state.
 36. The method of claim32 wherein said hydroprocessing catalyst comprises a plurality ofcatalytic particulates having a mean diameter ranging from about 35Tyler mesh to about 3 Tyler mesh; and a size distribution such that atleast about 90% by weight of said catalytic particulates have a diameterranging from R₁ to R₂, wherein:(1) R₁ has a value ranging from about1/64 inch to about 1/4 inch, (2) R₂ has a value ranging from about 1/64inch to about 1/4 inch, (3) a value of a ratio R₂ /R₁ ranges from about1.0 to about 1.4; andan aspect ratio of less than about 2.0.
 37. Themethod of claim 36 wherein said catalytic particulates have a sizedistribution such that at least about 95% by weight of said catalyticparticulates have a diameter ranging from R₁ to R₂.
 38. The method ofclaim 36 wherein said catalytic particulates have a size distributionsuch that at least about 97% by weight of said catalytic particulateshave a diameter ranging from R₁ to R₂.
 39. The method of claim 36wherein said catalytic particulates have a size distribution such that amaximum of about 2.0% by weight of said catalytic particulates have adiameter less than R₁.
 40. The method of claim 36 wherein said catalyticparticulates have a size distribution such that a maximum of about 1.0%by weight of said catalytic particulates have a diameter less than R₁.41. The method of claim 32 additionally comprising disposing a pluralityof inert pellets under said reaction zone prior to said introducing step(b).
 42. The method of claim 36 wherein each of said plurality ofannular mixture zones is essentially an annularly continuous mixturezone.
 43. The method of claim 42 wherein each of said plurality ofannular mixture zones has a generally uniform thickness.
 44. The methodof claim 36 wherein said catalytic particulates comprise aluminacarrying a metal oxide selected from the group consisting of molybdenumoxide, nickel oxide, and mixtures thereof.
 45. The method of claim 39wherein said catalytic particulates comprise alumina carrying a metaloxide selected from the group consisting of molybdenum oxide, nickeloxide, and mixtures thereof.
 46. The method of claim 44 wherein saidcatalytic particulates have a minimum mean crush strength of about 5lbs.
 47. The method of claim 46 wherein said aspect ratio is equal to orless than about 1.5.
 48. The method of claim 47 wherein said introducingstep (b) comprises upflowing through said catalyst of step (a) ahydrocarbon feed stream at a rate of flow such that said catalyticparticulates expand to less than 5% by length beyond a substantiallyfull axial length of the catalytic particulates in a packed bed state.49. The method of claim 47 wherein said introducing step (b) comprisesupflowing through said catalyst of step (a) a hydrocarbon feed stream ata rate of flow such that said catalytic particulates expand to less than1% by length beyond a substantially full axial length of the catalyticparticulates in a packed bed state.
 50. The method of claim 41additionally comprising disposing a plate member under saidhydroconversion reaction zone such that said plate member is totallyspaced therefrom, said plate member comprising a multiplicity of tubesbound thereto and axially extending downwardly therefrom for receivingsaid hydrocarbon feed stream and for conducting the same into a plenumchamber positioned under said hydroconversion reaction zone; anddisposing a permeable screen on said plate member for supporting saidinept pellets.
 51. A method for maximally occupying a reactor volumewith a substantially packed bed of hydroprocessing catalyst duringhydroprocessing by contacting the substantially packed bed ofhydroprocessing catalyst with an upflowing hydrocarbon feed streamhaving a liquid component and a hydrogen-containing gas componentcomprising the steps of:(a) disposing a substantially packed bed ofhydroprocessing catalyst in a reactor zone contained within a reactorvolume such that said substantially packed bed of hydroprocessingcatalyst occupies at least about 50% by volume of said reactor volume;(b) upflowing into said substantially packed bed of hydroprocessingcatalyst a hydroprocessing feed stream including a liquid component anda hydrogen-containing gas component and having a rate of flow such thatsaid substantially packed bed of hydroprocessing catalyst expands toless than 10% by length beyond a substantially full axial length of saidsubstantially packed bed of hydroprocessing catalyst in a packed bedstate; (c) withdrawing a volume of said hydroprocessing catalyst fromsaid reactor zone to commence essentially plug-flowing downwardly ofsaid substantially packed bed of hydroprocessing catalyst within saidreactor zone; and (d) adding hydroprocessing replacement catalyst tosaid essentially plug-flowing downwardly, substantially packed bed ofhydroprocessing catalyst of step (c) at a volume to substantiallyreplace said volume of said hydroprocessing catalyst of step (c). 52.The method of claim 51 additionally comprising repeating steps (b)-(d).53. The method of claim 51 additionally comprising forming, prior tosaid step (b) upflowing into said substantially packed bed ofhydroprocessing catalyst a hydroprocessing feed stream, a plurality ofannular mixture zones under said substantially packed bed ofhydroprocessing catalyst such that each of said annular mixture zonescontains said hydrocarbon feed stream and wherein said annular mixturezones are concentric with respect to each other and are coaxial withrespect to said reactor zone.
 54. The method of claim 51 wherein saidstep (b) upflowing into said substantially packed bed of hydroprocessingcatalyst a hydroprocessing feed stream comprises upflowing saidhydrocarbon feed stream from each of said annular mixture zones intosaid substantially packed bed of hydroprocessing catalyst.
 55. Themethod of claim 51 wherein said hydroprocessing catalyst comprises aplurality of catalytic particulates having a mean diameter ranging fromabout 35 Tyler mesh to about 3 Tyler mesh; and a size distribution suchthat at least about 90% by weight of said catalytic particulates have adiameter ranging from R₁ to R₂, wherein:(1) R₁ has a value ranging fromabout 1/64 inch to about 1/4 inch, (2) R₂ has a value ranging from about1/64 inch to about 1/4 inch, (3) a value of a ratio R₂ /R₁ ranges fromabout 1.0 to about 1.4; andan aspect ratio of less than about 2.0. 56.The method of claim 55 wherein said catalytic particulates have a sizedistribution such that at least about 95% by weight of said catalyticparticulates have a diameter ranging from R₁ to R₂.
 57. The method ofclaim 55 wherein said catalytic particulates have a size distributionsuch that at least about 97% by weight of said catalytic particulateshave a diameter ranging from R₁ to R₂.
 58. The method of claim 55wherein said catalytic particulates have a size distribution such that amaximum of about 2.0% by weight of said catalytic particulates have adiameter less than R₁.
 59. The method of claim 55 wherein said catalyticparticulates have a size distribution such that a maximum of about 1.0%by weight of said catalytic particulates have a diameter less than R₁.60. The method of claim 51 additionally comprising disposing a pluralityof inert pellets under said reactor zone prior to said upflowing step(b).
 61. The method of claim 53 wherein each of said plurality ofannular mixture zones is essentially an annularly continuous mixturezone.
 62. The method of claim 61 wherein each of said plurality ofannular mixture zones has a generally uniform thickness.
 63. The methodof claim 55 wherein said catalytic particulates comprise aluminacarrying a metal oxide selected from the group consisting of molybdenumoxide, nickel oxide, and mixtures thereof.
 64. The method of claim 58wherein said catalytic particulates comprise alumina carrying a metaloxide selected from the group consisting of molybdenum oxide, nickeloxide, and mixtures thereof.
 65. The method of claim 63 wherein saidcatalytic particulates have a minimum mean crush strength of about 5lbs.
 66. The method of claim 65 wherein said aspect ratio is equal to orless than about 1.5.
 67. The method of claim 66 wherein said upflowingstep (b) comprises upflowing through said catalyst of step (a) ahydroprocessing feed stream at a rate of flow such that said catalyticparticulates expand to less than 5% by length beyond a substantiallyfull axial length of the catalytic particulates in a packed bed state.68. The method of claim 66 wherein said upflowing step (b) comprisesupflowing through said catalyst of step (a) a hydroprocessing feedstream at a rate of flow such that said catalytic particulates expand toless than 1% by length beyond a substantially full axial length of thecatalytic particulates in a packed bed state.
 69. The method of claim 60additionally comprising disposing a plate member under said reactor zonesuch that said plate member is totally spaced therefrom, said platemember comprising a multiplicity of tubes bound thereto and axiallyextending downwardly therefrom for receiving said hydroprocessing feedstream and for conducting the same into a plenum chamber positionedunder said reactor zone; and disposing a permeable screen on said platemember for supporting said inert pellets.